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Bureau of Mines Information Circular/1984 




Review of Desliming Methods 
and Equipment 



By Christopher H. Roe 




UNITED STATES DEPARTMENT OF THE INTERIOR 



Information Circular 8972 



Review of Desliming Methods 
and Equipment 



By Christopher H. Roe 




UNITED STATES DEPARTMENT OF THE INTERIOR 
William P. Clark, Secretary 

BUREAU OF MINES 
Robert C. Norton, Director 









Library of Congress Cataloging in Publication Data: 



Roe, Christopher H 

Review of desliming methods and equipment. 

(Information circular ; 8972) 

Bibliography: p. 47-48. 

Supt. of Docs, no.: I 28.27:8972. 

1. Tailings (Metallurgy)— Dewatering. I. Title. II. Series: In- 
formation circular (United States. Bureau of Mines) ; 8972. 



TN2&5.U4- [TN535] 622s [622'.79l 83-600383 



CONTENTS 

Page 

Abstract 1 

Introduction 2 

Acknowledgments 2 

Chemical treatment • • • 2 

Advantages and disadvantages of chemical treatment 3 

Gravitational methods • 3 

Background ..........*. • 3 

Conventional thickeners 5 

High-rate thickeners 6 

Advantages and disadvantages of conventional and high-rate thickeners.. 7 

Multiple-plate thickeners 8 

Types of multiple-plate thickeners 10 

Advantages and disadvantages of multiple-plate thickeners 11 

Sizing thickeners 11 

Settling rate basis 11 

Solids flux basis 13 

Centrifugal sedimentation 15 

Background. 15 

Solid-bowl centrifuges 15 

Screen-bowl centrifuges 17 

Disk centrifuges 18 

Centrifuge performance 18 

Advantages and disadvantages of centrifuges 19 

Sizing centrifuges 20 

Sigma concept 20 

Pilot plant testing 21 

Filtration equipment 22 

Background 22 

Filter presses 22 

Advantages and disadvantages of filter presses 23 

Continuous pressure filters 24 

Belt filter presses 24 

Advantages and disadvantages of belt filter presses 26 

Vacuum filtration equipment 26 

Drum vacuum filters 26 

Cake discharge methods 27 

Equipment modifications 27 

Performance 28 

Advantages and disadvantages of drxim vacuum filters 29 

Rotary disk vacuum filters 29 

Advantages and disadvantages of rotary disk vacuum filters 30 

Horizontal continuous vacuum filters 30 

Rotary table vacuum filters..... «... 30 

Horizontal belt vacuum filters 31 

Advantages and disadvantages of horizontal vacuum filters 32 

Selecting and sizing filtration equipment 32 

Laboratory testing 33 

Factors affecting filtration 35 

Hydrocyclones 37 

Background 37 

Advantages and disadvantages of hydrocyclones 38 

Sizing hydrocyclones 38 



\ 



ii 



CONTENTS — Continued 



Page 



Thermal dewaterlng 39 

Background 39 

Thermal dryer operation 39 

Advantages and disadvantages of theimal dryers 40 

Sizing thermal dryers 41 

Current Bureau of Mines research on desllmlng methods 43 

Electroklnetlc methods 43 

Background 43 

Application 43 

Current research and use 44 

Trommel screen 44 

Description of the method and equipment 44 

Test results. 45 

Conclusions 45 

Discussion 45 

References 47 

Appendix A. — Mathematical terms 49 

Appendix B, — Manufacturers of dewaterlng equipment as of October 1982 54 

Appendix C, — Available dewaterlng equipment listed by manufacturer as of 

October 1982 57 

Appendix D. — Equipment efficiency 59 

ILLUSTRATIONS 

1. Cross section of gravity thickener showing zones of the sedimentation 

process 4 

2. Plan and cross-section views of conventional thickener 5 

3 . Cross section of high-rate thickener 7 

4. Diagram of particle path In ideal settling tank 8 

5 . Settling basin containing 10 parallel plates 8 

6. Settling basin containing 10 plates set at angle of 60° above horizontal.. 9 

7 . Cross section of countercurrent multiple-plate thickener 10 

8. Front and side cross sections of typical crossflow multiple-plate 

thickener 10 

9. Graph of A-B Interface height versus time for batch settling test 12 

10. Graph of total solids flux versus solids concentration showing relation- 

ships of the various components 14 

11. Graph of settling flux versus solids concentration showing operating line 

and limiting solids handling capacity •. 15 

12. Cross section of solid-bowl centrifuge 16 

13. Cross section of screen-bowl centrifuge 17 

14. Cross section of disk centrifuge..... 18 

15. Cross section of recessed-plate filter press 23 

16. Cross section of belt filter press 24 

17. Front and side views of drum vacuum filter 27 

18. Schematic drawings of various discharge methods for drum vacuum filters... 28 

19. Front and side views of disk vacuum filter.. 30 

20. Plan and cross-section views of horizontal rotary vacuum filter 30 

21. Cross section of typical horizontal belt vacuum filter 31 

22. Typical laboratory Installation for vacuum leaf tests 33 

23. Representative curve for cake formation rate versus feed solids 

concentration 35 



ill 



ILLUSTRATIONS—Continued 

Page 

24. Representative curve for cake formation rate versus cycle time 36 

25. Cross section of hydrocyclone 37 

26. Graph of pressure drop versus throughput capacity for two hydrocyclones . . 39 

27. Simplified cross sections showing operation of drum, suspension, and mul- 

t ilouver thermal dryers 40 

28. Simplified cross sections showing operation of vertical tray, continuous 

carrier , and f luidized-bed thermal dryers 40 

29. Two configurations for electrokinetically dewatering slimes 43 

30. Diagram showing operation of rotary trommel 45 

31. Chart showing generalized capabilities of commonly used dewatering equip- 

ment with respect to solids cake, moisture content, and particle size.. 46 
D-1. Example of a grade efficiency curve showing the relationship of X50, xgs, 

and Xmax 59 

TABLE 
1. Classification of selected vacuum filters 33 





UNIT OF MEASURE 


ABBREVIATIONS USED 


IN THIS REPORT 


°F 


degree Fahrenheit 


lb/in2 


pound per square inch 


ft 


foot 


min 


minute 


ft2 


square foot 


min/rev 


minute per revolution 


ftVh 


cubic foot per hour 


pet 


percent 


gal/h 


gallon per hour 


qt 


quart 


gal/min 


gallon per minute 


rev/min 


revolution per minute 


h 


hour 


ton/d 


ton per day 


in 


inch 


V/in2 


volt per square inch 


kW 


kilowatt 


wt pet 


weight percent 


lb 


pound 


yr 


year 


lb/h-ft-2 


pound per hour per 
square foot 







REVIEW OF DESLIMING METHODS AND EQUIPMENT 

By Christopher H, Roe 



ABSTRACT 

This Bureau of Mines report reviews the various methods of removing 
the liquid from tailings slurries that contain very fine grained solids. 
Gravitational settlement, centrifugation, filtration, and thermal drying 
methods are discussed in detail. Chemical additives, electrokinetic de- 
watering, and the rotary trommel screen are also mentioned. Lists of 
dewatering equipment and suppliers are given in the appendixes to assist 
the planner who must choose the most efficient, economic, and practical 
method of dewatering very fine grained mill tailings. 



^Mining engineer, Spokane Research Center, Bureau of Mines, Spokane, WA. 



INTRODUCTION 



The disposal of wet fine-grained wastes 
from a milling operation can be a diffi- 
cult challenge for the plant operator. 
He or she must consider the stability of 
the material after deposition, environ- 
mental consequences, and, above all, eco- 
nomic constraints. Waste material or 
tailings with a minimum of material 
smaller than 200 mesh (0.0029 in) can be 
piled up or used as backfill with few 
complications because it drains water 
easily and is relatively stable. 

On the other hand, materials with a 
high percentage of particles smaller than 
200 mesh present far more problems for 
water drainage and stability. The small 
size allows intermolecular attraction be- 
tween water molecules and solid particles 
to influence the hydraulic and structural 
characteristics of the material. As a 
result, when these fine tailings are sat- 
urated, they have low permeabilities and 
have little or no shear strength (16).^ 

Current tailings disposal practice is 
to mix these fine tailings or slimes (35, 
p. 1026) with large volumes of water and 
to pipe the resulting slurry to settling 
ponds. In these ponds, the solid parti- 
cles settle out of the slurry and the re- 
maining liquid is decanted out of the 
pond. This disposal method works well; 
however, it requires large ponds. Also, 
as a consequence of the small particle 



size, these slimes may not become com- 
pletely settled for years and may present 
a possibility of structural failure. 

These problems can be mitigated or 
eliminated if the water content is sub- 
stantially reduced prior to disposal. 
Many different methods have been devel- 
oped for separating small-sized solids 
from liquids, A cursory evaluation of 
these methods was done by the Bureau of 
Mines. This report presents the results 
of the study in a format that should as- 
sist mill operators who must establish a 
tailings disposal system or upgrade an 
existing dewatering circuit for metal or 
nonmetal mining operations; dewatering 
coal slimes is reviewed in a report done 
under contract to the Bureau of mines 
( 14 ) . Many of the conclusions in this 
report are based on experience from the 
coal processing, power generation, and 
sewage treatment industries in the hope 
that a sharing of knowledge will be bene- 
ficial to those interested in dewatering 
slimes. 

The descriptions of each dewatering 
method, along with the lists of manufac- 
turers and equipment in appendixes B and 
C, should assist mill operators in making 
decisions about the most appropriate and 
economic methods of tailings dewatering 
to investigate further. 



ACKNOWLEDGMENTS 



The author thanks Ken Miyoshi, 
mill superintendent, Western Nuclear 
Corp., Wellpinit, WA, for reviewing this 



paper and offering considerable technical 
advice. 



CHEMICAL TREATMENT 



This paper discusses the physical meth- 
ods of separating solids and liquids. It 
must be acknowledged, however, that any 
discussion of solid-liquid separation 
must include some comments on chemical 

^Underlined numbers in parentheses re- 
fer to items in the list of references 
preceding the appendixes. 



additives used in this process. Other 
Bureau projects are investigating the 
various aspects of chemical additives, so 
only a brief review of these substances 
is included in this report. 

Chemical treatment is often the first 
step in slurry concentration. Chemical 
additives (or reagents) are added to a 



slurry to promote formation of more eas- 
ily separated solid masses (24). For 
many of the dewatering devices evaluated 
in this report, chemical pretreatment 
is frequently used or may even be neces- 
sary to remove as much water from the 
solids as possible prior to mechanical 
dewatering. 

For many years, glue, gums, starch, 
lime, and similar additives were used as 
flocculating aids to improve the separa- 
tion rate of small solids from liquids 
(10). These substances were reasonably 
successful but increased the volume of 



the solids that had to be transported and 
disposed after dewatering (24). 

With the introduction of polymers or 
polyelectrolytes , great improvements oc- 
curred. Settling rates for solids in- 
creased by a factor of 10, 20, or more, 
and solids that could not previously be 
thickened were responsive to the addition 
of polymers (10). In addition, lower 
doses of these additives were needed to 
produce the desired thickening. As a re- 
sult, the volume of the solids did not 
increase and production rates improved 
(24). 



ADVANTAGES AND DISADVANTAGES OF CHEMICAL TREATMENT 



The proper application of chemical ad- 
ditives will — 

1. Increase separation efficiency. 

2. Increase throughput. 



3. Require a minimum 
mixing equipment. 



investment for 



There are, however, several disadvan- 
tages to using chemical additives — 



1. They can be very expensive. 

GRAVITATIONAL METHODS 



2. A large concentration of some addi- 
tives may be needed to produce the re- 
quired results. 

3. Not all slurries are responsive to 
the chemical additives (J_). 

Chemical additives are often needed to 
increase the efficiency or throughput of 
a dewatering system. The use of chemi- 
cal additives must be carefully consid- 
ered because they are expensive to pur- 
chase and will increase the processing 
cost. 



BACKGROUND 

This section describes equipment known 
as thickeners because of their capability 
to concentrate or thicken the solids from 
a slurry that has a very low solids 
content. In gravitational methods, the 
force of gravity causes the solids to 
settle and separate from the liquid. In 
the mining industry, flocculants are add- 
ed to the slurry which cause the solids 
to form larger masses or "floes" that 
settle at an accelerated rate. With 
respect to dewatering slimes , thickening 
equipment is often used to concentrate 
the fine particles for further processing 
by other dewatering equipment which will 
then produce the final dewatered product. 

Over the years, nomenclature has been 
adopted that is specific to thickener 



functions in the mineral processing in- 
dustries. The solid-liquid mixture that 
is to be separated by sedimentation is 
known as the feed. The sediment ed ma- 
terial having a solids content higher 
than that of the feed is the underflow. 
The clarified liquid from which solids 
have been removed is the overflow. This 
terminology is used even for equipment 
where the overflow does not migrate over 
a weir or the underflow does not emerge 
from the bottom of the sedimentation de- 
vice (J7). 

Regardless of name or variation in de- 
sign, all gravitational equipment depends 
on sedimentation to produce the thickened 
product. During initial sedimentation, 
solid particles in a nonturbulent solu- 
tion move downward under the influence of 
gravity relative to the liquid. The 



velocity of this movement increases until 
the upward force of drag, caused by the 
viscosity of the liquid, equals the down- 
ward force of gravity on the particles. 
These particles then fall at a constant 
velocity, called the terminal or free- 
settling velocity. In addition to 
gravity, the size of the reacting force 
is dependent on the particle diameter and 
density and the solution density and 
viscosity. The magnitude of the terminal 
settling velocity can be shown as — 



v= = 



_ Gx2(Ds-Di) 



J8, 



(1) 



where 



Vg = the terminal settling veloc- 
ity, feet per second, 

G = the gravitational constant, 
feet per second per second, 

X = the particle diameter, inch, 

Ds = the particle density, slugs 
per cubic foot, 

D| = the liquid density at a spe- 
cified temperature, slugs 
per cubic foot. 



boundary conditions such as the smallest 
particle size just completely retained by 
a given device ( 26 ) . 

As sedimentation continues, the concen- 
tration of the solids increases through 
the process of "hindered settling" and 
then continues into the "compression 
phase." In this phase, a further concen- 
tration of the sediment occurs and an in- 
crease in the concentration of solids 
takes place; however, the process pro- 
ceeds at a slower rate. This slowing 
down is due to the fact that during the 
exchange of solid matter for water, the 
water does not reach the top relatively 
unhindered but has to pass through in- 
creasingly narrowing capillaries between 
the more densely packed particles. Also, 
the friction between the touching parti- 
cles slows down the compression process 
(22). 

As the sedimentation process reaches 
equilibrixim in the tank, four zones will 
be present (fig. 1): 

1. Clear solution zone. 

2. Feed zone. 



and 



y = the viscosity of the liquid 
at a specified temperature, 
pound-seconds per square 
foot, 

J = a correction factor for par- 
ticle shape, dimensionless 
(26). 3 



Where nonspherical particles are con- 
cerned, Vs alters by factor (J), which is 
less than 1.0. Where there is a low den- 
sity of particles, Vg also decreases by a 
factor that is a function of the particle 
concentration. The purely mathematical 
description of sedimentation is imprecise 
owing to variations in particle diameter, 
shape, and distribution for any given 
slurry. Therefore, the above formula 
will be used for the determination of 



^A list of 
used in this 
dix A, 



all mathematical symbols 
report is given in appen- 



3. Transition zone. 

4. Compression zone ( 13 , p. 27-71). 

The clear solution zone contains the 
clarified solution of the overflow. The 
feed zone has the solids concentration of 
the unsedimented feed. The transition 
zone has a higher concentration of solids 
in hindered settling. The compression 

Feed 



Clear water 
zone 



Transition 
zone 




Compression 
zone 



FIGURE 1. - Cross section of gravity thickener 
showing zones of the sedimentation process (13), 



zone has the highest concentration of 
solids in compression and is the origin 
of the underflow. 

Gravity thickeners consist of open 
tanks with a feed inlet at the top and a 
means of collecting the sludge at the 
bottom by a rake. As the contents are 
slowly stirred by the rake rotation, the 
solids settle and are drawn off the bot- 
tom in a continuous underflow. Gases es- 
cape from the surface, and clear solution 
is removed by the overflow weir (24), 
Chemicals are added to the feed to aid 
settling, and, in the mining industry, 
many gravitational thickeners produce un- 
derflows having over 40 pet solids with 
recoveries approaching 100 pet. 

Three general types of thickeners will 
be discussed in this report: convention- 
al, high-rate, and multiple-plate thick- 
eners.. The following sections discuss 
each of these in detail. 

CONVENTIONAL THICKENERS 

Conventional thickeners use the sedi- 
mentation principle for liquid-solid sep- 
aration. They are much larger than high- 
rate or multiple-plate thickeners and 
have several characteristic design fea- 
tures such as — 

1. Cylindroconical shape. 

2. Annular overflow weir, 

3. Walkway and feed pipe support. 

4. Feed well. 

5. Drive mechanism and rake, 

6. Underflow cone or trench (17), 

Conventional devices are typified by a 
cylindrical upper portion attached to a 
shallow conical section having the apex 
oriented downward (fig, 2). The width or 
diameter of these thickeners is much 
larger than the depth. Most conventional 
thickeners are equipped with an annular 
overflow weir, which may be located in- 
ternal or external to the tank and which 




Overflow 



Feed well 




] — Walkway 



Annular 

overflow 

launder 



Walkway 



Rake 



• — U nde r f lo w 



Underflow 
pump 

FIGURE 2. - Plan and cross-section views of 
conventional thickener (17). 

is equipped with a froth baffle. Usual- 
ly, overflow is regulated through a 
notched weir so that adjustments can be 
made to compensate for uneven tank set- 
tlement in the subjacent soil. Most 
thickeners have a walkway to the center 
of the thickener. The walkway usually 
serves as support for the piping that 
carries the feed to the center of the 
machine. The feed pipe terminates at a 
device located at the center of the 
thickener, which is called the feed well 
or center well. The function of the feed 
well is to dissipate the kinetic energy 
of the incoming feed and to form a zone 
of quiescence conducive to sedimentation. 
Feed wells are manufactured in many 
shapes and sizes because recent infor- 
mation shows that the feed introduction 
method and the feed well shape greatly 
affect sedimentation behavior. The func- 
tion of the rakes in a thickener is to 
gently move the sedimented solids from 
the periphery towards the center 



discharge point. The movement of the 
rakes is provided by a drive mechanism. 
After the sediment ed solids have been 
moved toward the center of the tank by 
the rake, they are removed from the 
thickener through a cone or trench lo- 
cated near the bottom center ( 17 ) . 

Although they are all based on the same 
sedimentation process, a variety of con- 
ventional thickeners are being manufac- 
tured which differ from each other in 
their design. For instance, the feed 
pipe support may extend from the edge of 
the tank to the center or may span comr- 
pletely across the diameter of the tank. 
The rake arms in some thickeners are 
rigidly attached to a central vertical 
shaft or lattice, while in other designs, 
the arms may be suspended from cables. 
Drive mechanisms also vary from worm-spur 
gear combinations to hydraulically oper- 
ated push-pull arrangements for rotating 
the rake arms about the central axis. 
Each of these thickener configurations is 
designed to assist the process of con- 
tinuous sedimentation by steadily remov- 
ing the consolidated solids and the clear 
liquid to make room for the introduction 
of more feed material (17). 

HIGH-RATE THICKENERS 

Recently, several manufacturers have 
introduced high-capacity or high-rate 
thickeners , which have much smaller tanks 
than conventional thickeners. The lat- 
eral area for a conventional installation 
ranges from 5 to 10 ft^ for each ton of 
solids thickened per day, but the area of 
high-rate units may be as low as 0.3 to 
0.6 ft^ for each ton of solids per day 

These high-capacity thickeners have 
smaller tanks because they discharge the 
feed directly into the bed of settled 
solids and use chemical additives to 
hasten the flocculation of the sediments. 
The efficiency of the flocculation de- 
pends on how thoroughly the flocculant 
and the slurry are mixed. The many dif- 
ferent high-capacity thickeners on the 
market use various mechanisms or designs 
to mix the flocculant with the slurry. 



High-rate thickeners, while smaller 
than conventional models, can often meet 
or exceed the performance of the larger 
units. In one such comparison between a 
conventional and a high-rate thickener, 
the latter unit produced underflow solids 
concentrations equal to or greater than 
those produced by the conventional unit. 
This was done even at feed input rates 10 
times that of the conventional thickener. 
These results, though, were dependent on 
the slurry solids being responsive to the 
chemical flocculant used during the 
trials (10). 

Figura 3 shows a typical high-capacity 
thickener. The chemical additives are 
combined with the incoming feed and 
previously thickened solids in the mix- 
ing chamber. This mixture is thorough- 
ly churned by the blades of the mixing 
mechanism. The high concentration of 
the solids causes all chemical reactions 
to occur quickly and completely so that 
flocculation is considerably improved. 
This mixture moves from the reaction 
chamber to the clarified zone, where the 
flocculated solids quickly settle, to be 
used again with the incoming feed. A 
portion of the solids is collected inward 
by the rakes and evacuated as underflow. 
As in conventional thickeners, the clear 
solution migrates upward through the cir- 
culating sludge bed and exits over the 
overflow weir ( 23 , p. 19-52) . 

Gravity plays only a part in the 
solids-liquid separation in this thicken- 
er. It would be more precise to refer to 
this as a filtration unit in which the 
filter media is a suspended sludge bed. 
This is true because as the solids set- 
tle, the pore spaces between them become 
more restricted and trap other solids be- 
ing carried along with the water migrat- 
ing upward (17) . 

The horsepower required to operate 
these thickeners is approximately equiva- 
lent to the horsepower for conventional 
units having a similar capacity. The 
drive unit employed by high-rate thicken- 
ers is usually a hydraulic arrangement 
(17). 



Drive unit 



£l\|^^^Overf 



Feed inflow 
Thickened solids 

Rake arms 




Chemical additives 



low weir 



Mixing chamber 
Mixing mechanism 



-Underflow 

FIGURE 3. - Cross section of high-rate thickener (23). 



ADVANTAGES AND DISADVANTAGES OF 
CONVENTIONAL AND HIGH-RATE THICKENERS 

Thickeners have been used extensively 
by the minerals industry for concentrat- 
ing slurries, and over the years many 
design improvements have been made that 
enable them to be very reliable and effi- 
cient for thickening operations. Gravi- 
tational thickeners, in general, may be 
advantageous to an operation because 
they — 



1. Are 
operation. 



capable 



of 



continuous 



2. Are capable of processing slurries 
with a variety of solids concentrations 
and size ranges, 

3. Have fairly low maintenance and 
operational costs for the mechanical 
equipment. 

4. Provide a kneading action by the 
rake mechanism which is beneficial to the 
compression process (26), 



In addition, high-rate thickeners have 
some advantages over conventional thick- 
eners, such as — 

1, Smaller lateral space requirements 
(i.e, , lower installation costs), 

2. Greater throughput of solids based 
on available area for settlement (10). 

Gravitational thickeners, though, do have 
some drawbacks , which include — 

1, Requirements for large spaces, 

2, Necessity for large watertight 
basins. 

3. Sensitivity to persistent strong 
winds, especially for very large units. 

4. Expensive operation if large 
amounts of flocculants are needed to 
obtain the desired concentration (10). 



MULTIPLE-PLATE THICKENERS 

Innovations have been made to reduce 
the space needed for gravity thickeners. 
First among these is the multiple-plate 
thickener, which uses a series of evenly 
spaced inclined plates positioned in a 
settling tank. 

Multiple-plate thickeners are in use by 
the minerals industry around the world 
for clarifying, classifying, and thicken- 
ing. In the United States, this equip- 
ment is being used in the coal industry 
for dewatering the waste products of coal 
cleaning (7^). 

The multiple-plate thickener can be 
used to accomplish both free settling and 
hindered settling; however, for sim- 
plicity, only free settling will be dis- 
cussed in describing the theory of opera- 
tion. For an ideal settling basin, the 
thickener feed enters at one end of the 
basin, flows uniformly along its length 
at velocity V| , and exits at the other 
end (fig. 4). Any one particle will set- 
tle at velocity Vg. The actual trajec- 
tory of the particle is indicated by the 
vector Vy. If the trajectory takes the 
particle to the bottom of the basin be- 
fore it reaches the far end, then the 
particle is assumed to have been removed 
from the liquid. A particle starting at 




rea (Ap) 



Height 



idth 
KEY 

Ap Area of settling basin 
Qs Volumetric flow of incoming slurry 
V Flow velocity 
Vg Settling velocity 
Vy Particle velocity 



FIGURE 4. - Diagram of particle path in ideal 
settling tank (7). 

the top must settle to the bottom at 
velocity Vg in the same time or less than 
the velocity of the liquid, V| , in the 
basin. Thus, the feed quantity, Qg, di- 
vided by the settling area, Ap , of the 
basin is known as the overflow rate or 
surface loading and is often expressed as 
gallons per minute per square foot. 
Based on this relationship, all particles 
are removed that have a settling rate 
equal to or greater than the overflow 
rate. It should be noted that the height 
or detention time of the basin is not one 
of the main parameters that affect the 
separation efficiency (7). 



Area ( Ap) 



Feed rate—' 
( 10 QJ ■ 



KEY 
Ap Area of each plate 
Qs Volumetric flow of 
incoming slurry 

FIGURE 5. 



i ?^1^ Overflow 



■•>^-tV«.--f,|- 1llH-<lfcH[ III —■^■m« 



TT-i-iWie 




Total settling 
area = 1 Ox Ap 



Settled solids 



Settling basin containing 10 parallel plates (7). 



If the depth of the basin is reduced to 
a few inches and a number of such basins 
are stacked on top of each other, the re- 
sult is a simple multiple-plate sedimen- 
tation device. Figure 5 shows such a 
unit containing 10 plates, which can 
theoretically handle 10 times the flow 
rate as could the same basin without any 
plates. The liquid detention time is 
one-tenth as long, and the settling area 
is 10 times as large, so the same separa- 
tion efficiency is achieved because the 
overflow rate is unchanged (_7) . 

In practice, the plates are inclined so 
the settled particles will slide downward 
and the plates are, in effect, self- 
cleaning. Figure 6 shows an arrangement 
of 10 plates set at an angle of 60° above 
the horizontal. In this case, however, 
the plate area must be multiplied by the 
cosine of the angle to correctly deter- 
mine the capacity and overflow rate, be- 
cause only the projected area of each 
plate on a horizontal plane is counted. 
Thus, the total settling area is 10 times 
the plate area times cosine 60° , which 
equals 5Ap , and the capacity of the unit 
is 5Qs ij) . The following equation is 
used for determining the terminal set- 
tling velocity of the smallest grain size 
just completely retained: 



Settling area for 
one plate ( Ap) 



Ve = 



Sp V 



2 S 

Ln + E — cos Ai 

P sin 2 A, ' 



(2) 



where Vg = the terminal settling veloc- 
ity, feet per second, 

Sp = the spacing between plates, 
feet, 

V| = the feed flow velocity, feet 
per second, 

Lp = the plate length, feet, 

and A I = the inclination angle above 
horizontal of the plates , 
degrees (26). 



Overflow 




KEY 



A( Angle of plate 
inclination 



An Settling area 



Q3 Volumetric slurry 
flow 



Feed Inflow 
( 5Qs) 

FIGURE 6. - Settling basin containing 10 plates 
set at angle of 60° above horizontal (7). 

From equation 2, it appears that even 
at very high flow velocities the finest 
particles can be retained if the plates 
are long enough. This is not true, 
though, because the flow becomes turbu- 
lent with high flow velocities. At suf- 
ficiently high flow velocities, particles 
that may have been deposited on the 
plates will be flushed away when the 
force exerted on them by the flow becomes 
larger than the force of gravity on them 
(26). 

A typical thickener plate is usually 
2 ft wide and 10 ft long. Plate spac- 
ing is a critical variable because it 
must be large enough to prevent the 
settled solids from being disturbed 
by the upward flowing liquid but close 
enough to give the benefit of compact- 
ness. A 2-in spacing between plates 
is generally safe and is often used. 
The plate angle is another critical 
variable because the angle must be steep 
enough for the solids to flow or slide 
down the plate easily. For mineral- 
type solids having a high specific grav- 
ity, 45° is adequate; however, applica- 
tion experience is the best guide in this 
area (7). 



10 



TYPES OF MULTIPLE-PLATE THICKENERS 

Two different types of multiple-plate 
thickeners are available. Each type dif- 
fers according to the direction of the 
feed flow in relation to the inclination 
of the plates. The first type operates 
using the counterflow principle, where 
the feed material rises between the in- 
clined plates against the direction of 
inclination, as described in the previous 
section, to develop the principle of 
multiple-plate sedimentation. Figure 7 
shows a cross section of a typical coun- 
tercurrent thickener and its component 
parts. Note that the feed inflow line 
extends to the center of the tank in or- 
der to distribute the feed evenly to all 
plates (26). 

The second type of multiple-plate 
thickener uses the crossflow principle 
(fig. 8). The incoming slurry flows 
across the width of the inclined plates. 
The terminal settling velocity for the 
solids is determined from the equation — 



Vc = 



'p vi 



Lp cos A] 



(3) 



where the variables are the same as those 
defined for equation 2 (26). 



The crossflow multiple-plate thickener 
is more efficient in small-particle sep- 
aration and allows higher feed velocities 
than the counterflow model. This is be- 
cause the directions of the feed flow and 
the solids flow down the plate are per- 
pendicular and not opposed to each other 
as in the first type. There are practi- 
cal limits to the highest velocities 



Inclined 
plates 



Thickened 
solids 




rf lo w 



Feed inflow 



Underflow 

FIGURE 7. - Cross section of countercurrent 
multiple-plate thickener ( 26 ). 



FRONT VIEW 



SIDE VIEW 



Overflow 




Feed inflow 



Inclined plates 



Thickened solids 



Underflow 




FIGURE 8. - Front and side cross sections of typical crossflow multiple-plate thickener (26). 



11 



possible because the very fine solids may 
not have sufficient time to settle to the 
plate at very high feed velocities ( 26 ) . 

ADVANTAGES AND DISADVANTAGES OF 
MULTIPLE-PLATE THICKENERS 



and 



Wp(j = the weight ratio of water to 
solids in the discharge, 

Sf = the settling rate for a given 
Wps* feet per hour ( 13 , 
p. 27-72). 



The multiple-plate thickener is simple 
in design and operation. The only item 
requiring operator attention is the 
sludge withdrawal rake located at the un- 
derflow outlet. This is usually con- 
trolled by a variable-speed, positive 
displacement hopper, which is manually or 
automatically set by a detector monitor- 
ing the sludge level in the hopper (_7) . 

Spacing and inclination of the plates 
are very critical for an efficient op- 
eration; however, once these parameters 
have been determined, the multiple-plate 
thickener can reduce the space require- 
ments by as much as 90 pet compared with 
conventional thickeners (21). Multiple- 
plate thickeners are simple in con- 
struction and can be prefabricated for 
rapid erection (31) . These thickeners 
are normally constructed of mild steel 
or 3/16-in stainless steel. For corro- 
sive slurries, polyvinylchloride (PVC) or 
fiberglass-reinforced plastic plates can 
be used with rubber-lined tanks (_7). 

SIZING THICKENERS 

Settling Rate Basis 

The sizing of thickeners can be done 
using the method proposed by Coe and 
Clevenger. When the suspending fluid is 
water, the area necessary for particle 
settlement can be estimated by the fol- 
lowing formula: 



»^cs 



_ 1.333(Wrs - Wpd) 



(4) 



where 



•^cs 



the cross-sectional area of 
the basin, square feet per 
ton of dry solids thick- 
ened in 24 h. 



Wps = the weight ratio of water 
to solids in the slurry. 



This equation uses various values of 
Wps obtained from batch sedimentation 
tests. The procedure for these tests be- 
gins by mixing the solids to be tested in 
a graduated cylinder of water at a spe- 
cific water-to-solids ratio, W, 
cylinder is shaken, and the 
allowed to settle. The settling 



'rs» 
fines 



The 

are 

rate. 



Sf, is the subsidence rate for the solids 
in the bottom of the cylinder. Several 
values of Wpg should be tested, including 
the final discharge value, Wpd* The 
largest area, Acs» obtained will govern 
the size of the settling basin or number 
of settling plates (J^, p. 27-72). 

The capacity of a thickener can then be 
estimated by using the equation proposed 
by Deane. 



sc 



^ 1.33 ths (Sgs - 1) 



'gs 



(Sqsc Sqyy) 



(5) 



where 



Vgc = the volume of solids in com- 
pression, cubic feet per 
ton of solids per 24 h. 



ths = 



'gs 



the holding time for the 
solids to settle from the 
entering dilution of com- 
compression to the dilution 
of the discharge, hours, 

the average specific gravity 
of the solids. 



and 



Igy, = the specific gravity of wa- 
ter, which is 1.0, 



Sgsc - the average specific gravity 
of the solids in compres- 
sion (13, p. 27-72). 



The foregoing analysis assumes that 
the settling rate is dependent only on 
the feed solids concentration. Running 
settlement tests with different solids 



12 



concentrations will indicate the set- 
tling rates at different depths in a 
thickener tank. Talmadge and Fitch pro- 
posed another method based on the work 
done by Kynch, This method uses just 
one settlement test and provides an esti- 
mate of the coincident settling rate and 
solids concentration for various depths 
(32. p. 90). 

In this analysis, a slurry having the 
anticipated solids concentration of the 
full-scale operation is put into a tall 
cylinder. The height of the interface 
between zones A and B is recorded at fre- 
quent intervals. From these data, a 
graph of the interface height versus time 
is plotted as shown in figure 9 ( 18 , 
p. 12-32). 

The relationship among solids concen- 
tration, particle settling velocity, and 
interface height is shown by the follow- 
ing equation: 



E^ 



^SS HgbO 



(6) 



where C 



se 



"" Habe + Vs te ' 

= solids concentration at time 
te, pounds per cubic foot. 



and 



*^ss ~ solids concentration of the 
feed slurry, pounds per 
cubic foot, 

Ha bo ~ original height of A-B in- 
terface, feet, 

Habe = height of A-B interface at 
time te, feet, 

Vs = solids terminal settling ve- 
locity, feet per hour, 

te = an arbitrary time after 
solids settlement begins, 
hours (18, p. 12-36). 



The solids settling velocity, Vg, is 
the slope of the tangent to the curve for 
interface height versus time at time te. 
The quantity Hgbe + ^s ^e is the inter- 
cept of the tangent with the ordinate. 



:^ A 



m s^ 



I — Crit i 



Critical point 




30 60 90 120 150 180 210 240 
TIME, mjn 

KEY 

Height of A-B interface 

Tangent to curve 

A Zone A-clear solution 

B Zone B-solids at concentration of feed 

Zone C-transition between B and D 

D Zone D-solids in compression 

Habe Height of A-B interface at time tg 

Habo Original height of A-B interface 

tg Arbitrary time after settlement begins 

FIGURE 9. - Graph of A-B interface height ver- 
sus time for batch settling test (18). 

When the feed inflow rate, feed solids 
concentration, and underflow solids con- 
centration are established, the area of 
the thickener tank can be determined from 
the formula — 



'(^)fe-^)' ^'^ 



•^cs 



where A^s = tank area, square feet, 

Qs = volumetric feed rate, cu- 
bic feet per hour, 

and Csu = solids concentration of un- 
derflow, pounds per cubic 
foot (J^, p. 12-36; 32, 
p. 93). 

A series of solids concentration and 
settling velocity values can be deter- 
mined from the interface height versus 



13 



time curve. These values can then be 
used to establish and plot the solids 
concentration at different times. Sev- 
eral thickener areas are calculated using 
the corresponding values of Vg and Cge* 
A graph of A^s versus Vg can then be 
plotted. The maximum value for A^s vrLll 
be used for designing the 
pp. 12-36). 



thickener (18, 



Of the two methods discussed, the Kynch 
method is more often used for flocculated 
slurries. It must be remembered that 
these equations will provide only approx- 
imate values for sizing a conventional 
thickener. Pilot plant experimentation 
must be done using the slurries to be 
thickened in order to determine the opti- 
mum design features. 



After the required area has been estab- 
lished, the depth of the compression zone 
can be determined based on the retention 
time of the solids in this zone. This 
time is dependent on the rate of dis- 
charge and the concentration of the un- 
derflow. These values can be obtained 
from the circuit required for the under- 
flow solids concentration and from the 
interface versus time graph (32, p. 93). 

The volume and the height of the com- 
pression zone can then be computed: 



_ Qs c 



ss 



sc 



-<^)(l7^)- ^«' 



and 



where 



A Trd2 



- for a circular 
tank, (9) 



Vsc = volume of the compression 
zone, cubic feet, 

ths = retention time, minutes, 

Dg = density of the slurry 
solids, pounds per cubic 
foot, 

^ave ~ average slurry solids con- 
centration, pounds per cu- 
bic foot. 



and D| = density of the slurry liq- 
uid, pounds per cubic 
foot (32, p. 93). 



Solids Flux Basis 

The Coe-Clevenger and Kynch methods 
have been used many years for designing 
settling tanks; however, they have limi- 
tations because they are used to design a 
dynamic process based upon data obtained 
from a static settling test. A recent 
innovation for designing gravity set- 
tling equipment uses the solids flux 
concept. 

Solids flux is defined as the mass rate 
of solid flow through a unit area and 
portrays the conditions in a modern con- 
tinuous process gravitational thickener. 
The solids flux in a batch settling test 
is dependent only upon the velocity of 
the settling particles. In a continuous 
thickener, however, the solids flux is 
dependent not only on the settling ve- 
locity but also on the bulk transport 
of solids in the underflow. This can be 
represented as — 

Fts = Fbt + Fs = Css (vbt) + Css (vs) 



= Css (vbt + Vs), 



(10) 



where 



F-t-s = the total solids flux, 
pounds per square foot per 
hour. 



bt = 



the bulk transport flux 
component based on the un- 
derflow, pounds per square 
foot per h. 



The total height of the thickener tank 
can be determined by allowing 2-1/2 to 5 
ft above the compression zone for the 
solution in the clarified and transition 
zones (32, p. 94). 



Fg = the settling flux compo- 
nent, pounds per square 
foot per hour. 



14 



and 



Vbt = the bulk transport velocity 
of the solids, feet per 
hour, 

Vg = the settling velocity of the 
solids, feet per hour, 

Css = the solids concentration of 
the slurry, pounds per cu- 
bic foot (32, p. 95). 



Figure 10 shows a graphic relationship 
among the various components of the 
solids flux for a continuous gravity 
thickener. This graph can be constructed 
for any gravity thickener, once the un- 
derflow rate and the bulk transport 
velocities have been determined ( 32 , 
p. 96). 

The figure indicates that for low 
solids concentrations, most of the total 
flux is composed of the settling comr- 
ponent. As the solids concentration is 
increased, the bulk transport component 
represents a correspondingly greater por- 
tion of the total flux. When the total 
flux is less than the solids handling 
capacity, there is only one associated 
value of solids concentration, and there 
will be a zone of constant concentration 
for any particular underflow rate. If 
the total flux is greater than the solids 
handling capacity, there will be two 
zones of constant concentration for any 
particular underflow rate, but the solids 
handling capacity represents the maximum 
practical value for stable thickener op- 
eration (29). 



Mathematical manipulation of 
ables indicates that 



Fs = Fts 



i-t)' 



the vari- 



(11) 



where 



Csu = the solids concentration of 
the underflow, pounds per 
cubic foot. 



When designing a thickener, the total 
solids flux, Fts» and the underflow 
solids concentration, Csu> will be known 
or fixed, but the settling flux comr- 
ponent, Fg, and solids concentration. 



X 

3 



(0 

o 

_l 

o 
m 

_i 
< 
1- 
o 



Batch flux curve 




Bulk transport 



SOLIDS CONCENTRATION. Css 

FIGURE 10. - Graph of total solids flux versus 
solids concentration showing relationships of the 
various components (32). 



'ss » 



will be unknown. Using values for 
F-j-s and Cgs ^^^ picking values for Cgu* 
equation 11 can be used to obtain cor- 
responding values for Fg. A graph of the 
settling flux, Fg, versus concentration, 
Css» ^^^ then be constructed, as shown in 
figure 10 (32, p. 97). 

If values of Cgs and the corresponding 
Fg are plotted as shown in figure 11, the 
result is a straight line with a slope of 
-Fg/Csu and an intercept of Fg i . This is 
referred to as the operating line and 
represents the thickening characteristics 
for a particular thickener at any speci- 
fied underflow rate. The batch flux 
curve can also be plotted, and the inter- 
sections of the operating line and the 
batch flux line indicate those solids 
concentrations that satisfy both rela- 
tionships (32, p. 97). 

In figure 11, an operating line has 
been plotted for a thickener that is 
functioning within its limits. In this 
example, only one zone of concentration 
exists at Cgsi to produce an underflow 
concentration, Cgui* If» however, the 
slurry concentration and the underflow 
concentration are increased to Css2 and 
Csu2» respectively, the operating line 
intercepts the tangent to the batch flux 
curve and a second zone of concentration 



15 




-SS1 ^SS2 



SOLIDS CONCENTRATION, C35 

FIGURE 11. - Graph of settling flux versus sol- 
ids concentration showing operating line and lim- 
iting solids handling capacity (32). 

begins to appear with a concentration of 
Cs2« This condition represents the lim- 
iting solids handling capacity of a 
thickener at a particular underflow rate 
(32, p. 97). 

In practice, the necessary data are ob- 
tained by running tests using a thickener 
and the slurry to be thickened. Altering 
the applied flux, Fts» ^^^ the underflow 



concentration, v^su> 



dlcatlon of 



the 



will provide 
limiting flux 



After the maximum value of C 



BACKGROUND 



s u 



an in- 
value . 
has been 



determined, a tangent to the batch flux 
curve can be drawn. The area of the 
tank can then be determined by dividing 
the maximum rate of solids loading by 
the limiting flux. Experience has 
shown, though, that the optimal thicken- 
er throughput will be about 90 pet of 
the values calculated ( 32 , p. 98). 

As with conventional thickeners , the 
use of a pilot plant is also necessary 
for determining the best size for a high- 
capacity thickener and for establishing 
the amount of flocculant needed for 
proper thickening. The properties of the 
slurry and the many available flocculants 
will vary considerably, so the best flow 
rates and flocculant injections should be 
determined experimentally. 

Multiple-plate thickeners should also 
be sized according to results of labora- 
tory and pilot plant testing. The re- 
sults will aid the designer in determin- 
ing the settling rate of the solids so 
that the overflow rate and effluent qual- 
ity can be established. The sludge vol- 
ume can also be determined in order to 
establish the underflow solids concentra- 
tion. Finally, the need for chemical 
pretreatment can be evaluated (7). 



CENTRIFUGAL SEDIMENTATION 



Centrifugal sedimentation depends on 
the density difference between solids and 
liquids where the particles are subjected 
to centrifugal forces that make them move 
radially outwards or Inwards through the 
liquid, depending on whether they are 
heavier or lighter than the liquid ( 32 , 
p. 125). Centrifuges are compact ma- 
chines and are capable of producing high 
liquid clarification and solids concen- 
tration. Most of the units available 
today are designed for continuous opera- 
tion (24). 

Centrifuges have relatively modest cap- 
ital costs, but they may be expensive to 
operate because of the need for chemical 
conditioning in most applications, high 
power consumption, and extensive mainte- 
nance requirements. Lower speed units 



consume less power and, as a result, have 
fewer wear problems. Maintenance diffi- 
culties can be greatly reduced if the 
construction materials are specified to 
match the abrasive or corrosive charac- 
teristics of the slurry to be handled 
(24). 

A large variety of centrifuges are 
available on the market; however, several 
types are used primarily for clarifying 
and not dewaterlng, so they are not in- 
cluded in this report. Solid-bowl, 
screen-bowl and disk centrifuges are of 
interest in dewaterlng slimes and are 
discussed in the following sections. 

SOLID-BOWL CENTRIFUGES 

Solid-bowl, scroll, or decanter cen- 
trifuges consist of a horizontally rotat- 
ing chamber that has one end tapered into 



16 



a cone. The slurry is admitted through 
axial feed tubes and removed radially out 
of the bowl, A screw or scroll mechanism 
rotates in the same direction as the bowl 
but 5 to 100 rev/min faster or slower 
than the bowl and thus can push the sol- 
ids along the length of the chamber. The 
speed of the bowl rotation can vary from 
1,600 to 6,000 rev/min. The solids are 
collected by the scroll towards the 
tapered end of the bowl, while the solu- 
tion overflows a weir at the other end. 
The tapered section serves as a drying 
zone or beach area prior to the cake dis- 
charge ( 24 , 32 , p, 139), Figure 12 shows 
a simplified cross section of a typical 
solid-bowl centrifuge. 

With regard to the machine design, a 
number of variations are available in the 
contour of the centrifuge shell, scroll 
flight angle and pitch, beach angle and 
length, conveyor speed, and feed posi- 
tion. An alternative to the liquid over- 
flow outlet is an internally mounted tube 
for skimming off the liquid (^2, p, 140) . 

Specially designed models of this cen- 
trifuge have been used to dewater very 
fine slurries in hydrocyclone circuits 
for recovering and dewatering deslimed 
coal. Other solid-bowl centrifuges have 
been used to dewater froth flotation 
tailings after thickening in a static 
thickener (18), Another major area of 
application for this type of centrifuge 
is in the classification of solids such 
as kaolin clay and titanium oxide ( 32 , 
p. 140), 

Polyelectrolytes are widely used for 
the flocculation of the solids to be 

Rotating solid bowl 

Rotating scroll 

^ / /J- Drive sheave 

Gear unit 




Liquid 
discharge 



FIGURE 12, - Cross section of solid- bowl centrifuge (3), 



dewatered in solid-bowl centrifuges. The 
point of addition varies , depending on 
the type of the polyelectrolyte and slur- 
ry. Anionic polyelectrolytes are usually 
introduced upstream from the centrifuge, 
while cationics are added within the cen- 
trifuge because they react very quickly 
with the slurry ( 32 , p, 140), 

Wilson and Miller ( 36 ) compared the ef- 
fectiveness of solid-bowl centrifuges to 
that of disk vacuum filters and commented 
on the parameters that affect their use 
in coal dewatering. The solid-bowl cen- 
trifuge and the disk vacuum filter pro- 
duced similar results on the same slurry. 
The solids from each of the two dewater- 
ing machines contained 20 pet minus 325 
mesh with 20 pet surface moisture and a 
centrate or filtrate of about 0,1 to 
0,2 pet solids. The products from these 
units are very similar in nature, so 
the choice between the devices should be 
made on the basis of handling character- 
istics, throughput, floor space, cost, 
and overall circuit or layout considera- 
tions. If the centrifuge is chosen, in- 
plant adjustments to feed rates, speed, 
and pool depth can be made to obtain the 
proper balance of throughput, moisture, 
maintenance cost, and effluent clarity 
(36), 

Wilson and Miller also found that 
steam-heating the feed to 110° F resulted 
in a 4-pct reduction in product moisture; 
however, higher temperatures did not fur- 
ther improve results. The use of heat 
may be economically feasible as a means 
of reducing the product moisture if an 
inexpensive source of heat is available 
(36), 

One of the principal advantages of this 
machine is that it can dewater dilute 
slurries. In plants that use shaking ta- 
bles , no dewatering screens are required 
between the table and the centrifuge as 
would be needed when a screen-type ma- 
chine is used. The solid-bowl unit will 
require more horsepower, though, because 
it must accelerate the water as well as 
the solids during the dewatering opera- 
tion (18, p, 12-20), 



17 



SCREEN-BOWL CENTRIFUGES 

Screen bowls , also called basket or 
perforate units, are a second type of 
centrifuge. Positive-discharge machines 
are screen-bowl centrifuges with trans- 
port devices and are the most common 
type of centrifuge found in the minerals 
industry today. These units have two 
elements that rotate about a vertical 
or horizontal axis. These elements con- 
sist of an outside conical screen frame 
and an inside solid cone that carries 
spiral hindrance flights. A gear ar- 
rangement produces a differential speed 
in the two rotating elements so that they 
both rotate in the same direction, but 
the screen element moves slightly faster 
than the cone carrying the spirals. 
The operation is similar to that of the 
solid-bowl centrifuge. The slurry en- 
ters the machine at the top and falls on 
the apex of the cone. The centrifugal 
force developed by the rotating cone 
throws the solid-liquid mixture against 
the screen. The water passes through the 
perforations and is collected in an 
effluent chamber. Meanwhile, the flights 
spiral downward and the solid material 
is gradually transported to the bottom 
of the screen. The conical shape of 
the basket causes the solids and water 
to be subjected to zones of increasing 
centrifugal force ( 18 , p. 12-16). Fig- 
ure 13 shows a cross section of a typical 
screen-bowl centrifuge. 

Another type of screen-bowl centrifuge 
is the vibrating-basket type that is fre- 
quently being installed in new plants. 
This centrifuge has either a vertical or 
horizontal basket that vibrates in such a 
manner as to cause the solids to move 
through the machine. This vibration 
tends to loosen the bed of particles so 
that they are free-draining and only mod- 
erate force is required to effect thor- 
ough dewatering. Because of the low 
speed generally used in these centri- 
fuges, the moisture content of the solids 
is usually higher than that produced by 
the transport-type unit; however, wear 
and horsepower are low and solids 
degradation is minimal. The principal 
difference between the horizontal and 



Feea 




Rotating screen 
bowl 



Rotating cone with 
spiral flights 



Liquid discharge 



Screen bowl 
support arms 



Solids discharge 



Gear drive 
mechanism 



FIGURE 13, - Cross section of screen-bowl 
centrifuge (5), 

vertical screen-bowl types is that the 
horizontal axis machine requires less 
headroom than the vertical device ( 18 , 
pp. 12-16 to 12-19), 

Wilson and Miller also conducted tests 
on screen-bowl units. They found that the 
positive-discharge screen-bowl centrifuge 
provided 4 to 6 pet lower moisture in the 
solids product than the solid-bowl cen- 
trifuge. Thus, for similar sized centri- 
fuges , the screen bowl is preferable to 
the solid bowl for this particular in- 
stance. As with the solid-bowl centri- 
fuge, they found that the screen-bowl 
centrate imist be bled out of the plant to 
a pond or backed up by a secondary recov- 
ery system. Because of the recovery re- 
quirement, this type of centrifuge should 
operate for maximum moisture reduction as 
opposed to maximum effluent clarity to 
fully benefit from this costly system 
(36). 

Further test results indicated that on- 
ly minor moisture reduction was obtained 
by steam heating the feed, so heating was 
not recommended for this type of centri- 
fuge. Surface-tension-reducing chemicals 
and a flocculant were also tested; how- 
ever, both approaches were ineffective in 
reducing moisture (36). 



18 



DISK CENTRIFUGES 

The last type of centrifuge to be dis- 
cussed is the disk centrifuge. Its con- 
struction is similar to that of the ver- 
tical screen-bowl unit; however, instead 
of using just one cone, it uses multi- 
ple cones for dewatering (fig. 14). The 
basic idea of increasing the settling 
capacity by using a number of disks in 
parallel is the same as the multiple- 
plate principle in gravity sedimentation 
(312, p. 141). 

The slurry enters the unit at the top, 
and the solids are forced to the outer 
circumference by centrifugal action, then 
removed through the outer rim. The clar- 
ified water is channeled upward through 
passages between the disks. These cen- 
trifuges can handle inflows of up to 
3,300 gal/min containing low-density par- 
ticles up to 0.1 in in diameter and con- 
centrations to about 1 pet solids. The 
output can be up to 6 pet solids , or even 
up to 10 pet solids if chemical additives 
are used (24). 

Disk centrifuges are operated at speeds 
up to 12,000 rev/min, depending on the 
bowl diameter. The bowls usually have 

Feed 



rp-^ Liquid 

Rotating disks 




Solids 



Rotating 
screen bowl 



FIGURE 14, - Cross section of disk centrifuge (32). 



equal dimensions of height and diameter 
for optimum capacities, and the angle of 
the cones is usually between 35° and 50°, 
which is large enough to facilitate the 
sliding of the particles on the disk sur- 
faces ( 32 , p. 143). 

There are several variations in this 
design, which include recirculating the 
solids discharge, a facility for wash- 
ing before discharge, and a paring tube 
for pressurized solids discharge ( 32 , 
pp. 143-145). 

Disk centrifuges are very effective for 
dewatering fine-grained solids and are 
often used for dewatering kaolin clay 
(22, p. 144). 

CENTRIFUGE PERFORMANCE 

The performance and efficiency of a 
centrifuge depend on a number of factors. 
The more important factors are — 

1. Centrifuge rotation speed. 

2. Diameter. 

3. Length. 

4. Beach angle and length (for hori- 

zontal centrifuges) . 

5. Feed point of slurry. 



6. 
7. 

8. 



Feed point of flocculants. 

Scroll rotation speed differential 
and pitch (for positive-discharge 
units) . 



Pool depth (for 
(11, p. 4-3). 



solid-bowl units) 



Increasing the bowl rotation speed 
usually increases the solids recovery. 
There may be an increase in the solids 
cake concentration; however, the increase 
of the fines in the cake tends to in- 
crease the cake moisture. Higher speeds 
also increase both maintenance and op- 
erating costs of the centrifuge ( 11 , 
p. 4-3). 



19 



Centrifuge-bowl diameters generally 
range from 6 to 50 in, and bowl lengths 
are generally from two to four times the 
bowl diameter. Bowl speed is normally a 
function of the bowl diameter because the 
effects of speed and diameter determine 
the resulting centrifugal force acting on 
the slurry. Typical values for cen- 
trifugal force range from 1,000 to 4,000 
times the force of gravity; the higher 
centrifugal forces are associated with 
the smaller bowl diameters ( 11 , p. 4-3). 

The bowl length will affect the cen- 
trate clarity, A longer bowl increases 
the residence time of the slurry as it 
travels from one end of the bowl to the 
other. This increased time allows the 
finer particles sufficient time to sep- 
arate from the liquid (2^, p. 11-35). 

In solid-bowl units, the beach angle 
and beach length will affect both the 
percent solids in the final cake and the 
torque needed to move the solids out of 
the centrifuge at a constant scroll dif- 
ferential speed. The longer the cake is 
allowed to dewater on the beach, the 
higher the cake solid concentration, and 
the higher the torque requirements for 
discharge ( 11 , p. 4-4) , 

The feed entry point into the bowl 
will influence the percent solids in the 
cake and the solids recovery. For solid- 
bowl centrifuges , recovery will be im- 
proved if the feed entry point is near 
the beach because the slurry particles 
have a longer distance to migrate to the 
end of the machine where the liquid exits 
(JJ., p, 4-4), 

Chemicals may be added to the slurry to 
accelerate the flocculation of fine par- 
ticles that do not immediately separate 
from the liquid. The high degree of agi- 
tation and mixing within the centrifuge 
generally necessitates large polymer 
doses to effect an increase in the solids 
recovery. In one instance, a dosage 
range of 0,1 to 0,3 lb polymer per ton of 
solids increased the capacity of a cen- 
trifuge 5 to 10 pet at the same solids 
recovery (11, p, 4-9), 



In positive-discharge units , changes in 
the scroll rotation speed and pitch will 
affect the solids recovery and cake dry- 
ness. Small residence times allow only 
the heavier and larger solids to be re- 
moved, while the finer solids remain sus- 
pended in the liquid. This factor will 
be particularly important when dewatering 
slimes. 

Most solid-bowl centrifuges are provid- 
ed with an adjustable pool depth setting 
so that the liquid level in the bowl can 
be changed after Installation. An in- 
crease in the pool depth will improve the 
percentage of the solids that are sep- 
arated from the liquid. This improvement 
in solids recovery is due to an increased 
residence time and a reduction in the 
agitation within the centrifuge at deeper 
pool depths; however, deeper ponds will 
increase the amount of moisture in the 
cake because of the reduced dewatering 
time on the beach (11, p, 4-4), 

ADVANTAGES AND DISADVANTAGES 
OF CENTRIFUGES 

A main advantage of the centrifuge is 
its operational flexibility. Within its 
design limits, a centrifuge can be fed a 
slurry at various rates and still provide 
a consistent solids product. If the feed 
rate exceeds the design limits of the 
unit, the excess solids appear in the 
liquid; however, the quality of the dis- 
charged cake does not deteriorate and the 
percent solids remains relatively con- 
stant (JJ_, p. 4-4), 

A disadvantage of the centrifuge is 
abrasive wear on the scroll and other 
interior parts , which results in high 
maintenance costs. In recent years, 
scrolls have been manufactured with im- 
proved materials such as tungsten carbide 
on the wearing surfaces. Operational 
results indicate an order of magnitude 
improvement in the life of these com- 
ponents (3; JJ_, p, 4-4), Another draw- 
back of the centrifuge is that the feed 
slurry may need to be prethickened. 
Although centrifuges can dewater a wide 
range of slurries , a very low feed solids 



20 



content means that the centrifuge must 
process large volumes of slurry input. 
The number of centrifuges needed in a de- 
watering circuit is directly proportional 
to the volume of feed slurry to be pro- 
cessed. Consequently, the slurry should 
be prethickened by sedimentation, hydro- 
cycloning, or other means prior to cen- 
trifugation. The final circuit configu- 
ration will be the result of compromising 
the performance and economy of the pre- 
thickening and centrifuge equipment ( 11 , 
p. 4-7). 

SIZING CENTRIFUGES 

When centrifuging a slurry, it is not 
necessarily true that an increased force 
will decrease the moisture content of 
the product. Materials that deform, 
break, or degrade will not be dewatered 
proportionally to the applied forces. 
It should also be realized that horse- 
power, wear, maintenance, and degrada- 
tion will accelerate with increased 
forces applied in the machine. While a 
large centrifugal force developed by a 
machine may be an indication that it is 
sturdily built, this force should not be 
the only criterion used in selecting a 
centrifuge to dewater any material ( 18 , 
p. 12-15). 

For any particle traveling in a circu- 
lar motion about a point, the centrifugal 
acceleration is 

Ac =^, (12) 



Centrifugal acceleration is then ex- 
pressed as multiples of the gravitational 
acceleration: 



and 



Ac ^ Vp2 
G GR^' 



(13) 



p ^ Vp2^ (2TrRcN)2 ^ Att^R^N^ 
^ GRc GRc G * 



where Vp = 2 RcN, feet per second, 

N = the number of revolutions per 
second, 

TT = 3.1416, 

Fc = the centrifugal acceleration 
expressed as multiples of 
gravitational acceleration, 
dimensionless , 

and G = the gravitational accelera- 
tion, feet per second per 
second (J^, p. 12-14). 

SIGMA CONCEPT 

The "sigma concept" has been widely 
used in the field of centrifugal sedi- 
mentation for the last 30 yr. It is a 
simplified representation of machine 
performance in terms of the particle 
size, the total volumetric rate, and an 
index of the centrifuge size. The sigma 
concept characterizes a centrifuge's 
ability to separate solids from liquids 
and is widely used in industry ( 32 , 
p. 130). 



where Ac = centrifugal acceleration, 
feet per second per 
second. 

Vp = the linear peripheral ve- 
locity, feet per second, 

and Re = the radius of curvature, 
feet (18, p. 12-14). 



The volumetric throughput of a centri- 
fuge can be expressed as — 



Qc = 2 V. 



27rL3zf(|Rcb2 +^Rsz2)], (15) 



where Qc = the volumetric throughput, 
cubic feet per second. 



21 



Vs = the terminal settling veloc- 
ity of the solids, feet per 
second, 

0) = the angular velocity of the 
solids, radians per second, 

G = the acceleration of gravity, 
feet per second per second, 

jsz ~ the length of the settling 
zone, feet. 



because different total efficiencies can 
be obtained for a given cut size, depend- 
ing on the size distribution of the sol- 
ids. The best method of describing the 
performance of a centrifuge is by using 
the grade efficiency curve, which is 
briefly discussed in appendix D. This 
requires many tests, together with deeper 
theoretical considerations, but the re- 
sults will provide more meaningful and 
reliable predictions of total efficien- 
cies for different slurries (32, p. 133). 



and 



Rsz = the radial length from the 
rotational axis to the set- 
tling zone surface, feet. 



s-cb 



the radial length from the 
rotational axis to the cen- 
trifuge bowl surface, feet 
(11, p. 4-5). 



This equation is composed of two parts. 
The first component, Vg , describes the 
settling characteristics of the solids. 
The remainder of the equation represents 
the machine variables that effect separa- 
tion efficiency, such that — 



2 = 2itLc 






R 



sz 



')' 



(16) 



where Z = the sigma value for a par- 
ticular centrifuge configu- 
ration (11, p. 4-5). 



PILOT-PLANT TESTING 

The centrifuge evaluation procedure, 
generally used in pilot testing, relates 
the percent of solids recovered and cake 
solids concentration to the feed rate 
at various operating conditions. The 
recovery of feed solids in the centri- 
fuge cake is determined from the mea- 
surements of flow rates and concentra- 
tions entering and leaving the centri- 
fuge. By combining the material balance 
equations and the definition of recov- 
ery, the following simplified equation 
relates the recovery to various solids 
concentrations : 



R 



sc 



_ "sc 



(Wss - Wsi) 



where R 



Wss (W 



sc 



Wsl)* 



(17) 



sc = the percent recovery of 
solids in the cake, 



Equation 16 is the basic expression of 
the sigma concept, which gives an esti- 
mate of the maximum flow rate that will 
allow solids of a particular size to sep- 
arate from the liquid. Sigma is a con- 
stant containing factors pertaining only 
to the centrifuge and can be thought of 
as the theoretical capacity factor. It 
is expressed in terms of area and facili- 
tates comparison between the performances 
of geometrically and hydrodynamically 
similar centrifuges processing the same 
slurry (32, p. 131). 

A shortcoming of the sigma concept is 
that the cut size or smallest particle 
separated from the liquid is not suitable 
as a criterion for separation efficiency 



Wsc = the weight percent of sol- 
ids in the cake, 

= the weight percent of sol- 
ids in the centrate. 



W 



s I 



and Wgs = the weight percent of sol- 
ids in the fed slurry ( 11 , 
pp. 4-6 to 4-7). 

This equation simplifies the testing pro- 
cedure so that the solids recovery can be 
calculated from concentration measure- 
ments made on the three process streams 
entering or leaving the centrifuge. 
Thus, flow rate measurements need not be 
made to calculate the recovery of solids 
in the cake (11, p. 4-7). 



22 



FILTRATION EQUIPMENT 



BACKGROUND 

Filtration is a process where solids 
are separated from liquids by passing a 
slurry through a permeable filter medium 
that retains the solids. To cause the 
fluid to flow through the filter medium, 
a pressure drop has to be applied. The 
pressure drop can be achieved by various 
means, including gravity, vacuum, or di- 
rect pressure (32, p. 171). If the liq- 
uid is induced to flow through the medium 
by hydrostatic head, it is called gravity 
filtration. If higher than atmospheric 
pressure is applied upstream from the 
filter, the process is called pressure 
filtration; if lower than atmospheric 
pressure is applied downstream from the 
filter, it is referred to as vacuum fil- 
tration (20, p. 1473). 

There are two types of filtration sys- 
tems available: surface filters and 
depth filters. Surface filters are used 
for cake filtration, where the solids 
are deposited in the form of a cake on 
the surface of a relatively thin filter 
medium. Depth filters are used for deep- 
bed filtration, where particle deposi- 
tion takes place within the medium (32, 
p. 173). Surface filtration is the pro- 
cess in common use by the minerals indus- 
try and is the only one discussed in this 
report. 

At the commencement of surface filtra- 
tion, the slurry particles that are the 
same size as or larger than the openings 
of the filter medium are held at these 
openings and create smaller passages , 
which remove even smaller particles from 
the slurry. A filter cake is formed 
which, in turn, functions as a more effi- 
cient filter for the subsequent filtra- 
tion (32, p. 172). 



filter equipment is usually expressed as 
the yield in pounds of dry solids per 
square foot of filter area per hour (24). 

The filter medium is probably the sin- 
gle most important component of a filter 
that affects equipment performance. Many 
different materials are now available as 
filter media. Filter cloth can be com- 
posed of duck, chain, and twill weaves, 
as well as felt made of cotton, ny- 
lon, polyester, polypropylene, and other 
natural or synthetic materials. Filter 
cloths and screens made of steel, stain- 
less steel, and other alloys are also on 
the market (29). 

Filter paper is often required as the 
filter medium to obtain high solids 
retention. Filter paper requires the 
support of filter cloth, screens, or per- 
forated metal sheets to prevent its 
breaking. The filter paper can be easily 
removed from the support filter medium 
(29). 

Covering the surface of the filter 
cloth or filter paper with a filter aid 
is known as a precoat and is often re- 
quired to prevent blinding or clogging 
the filter medium. Materials such as di- 
atomaceous earth, paper pulp, or perlite 
can be very effective. A filter aid can 
also be added to the slurry as a body 
feed to produce a filter cake that re- 
mains relatively permeable during the 
filtration cycle and gives a good overall 
filtration rate (29). Various combina- 
tions of filter cloth, filter paper, and 
precoating are capable of filtering out 
particles as small as 0.0002 in in diame- 
ter (3a). 

The following sections discuss the fil- 
tration equipment commonly used today. 



In practice, solids are deposited on 
the filter medium as a cake, which is re- 
moved from the filter medium by a mechan- 
ical method such as scraping. After the 
cake is removed, the filter is cleaned by 
spray washing, then put into position to 
receive more solids. The performance of 



FILTER PRESSES 

Filter presses are batch units that use 
higher than atmospheric pressures and a 
series of filter plates to separate the 
solids from the liquid. They are also 
called pressure, plate-and-f rame, or 



23 



press filters. Although this type of de- 
vice is well known, it has had little at- 
tention or application in the U.S. miner- 
als industry because of its batch process 
nature, high capital cost, and associated 
high labor requirements. In spite of 
these drawbacks, the filter press should 
receive close attention because of its 
separation efficiency (24). 

There are two types of filter press 
systems available — the high-pressure sys- 
tem and the low-pressure system. The 
first type uses a pressure of up to 225 
Ib/in^ , while the second type uses a max- 
imum pressure of 100 Ib/in^ . Research 
has shown that the higher filter pressure 
does not provide a significant benefit 
over the lower pressure unit with respect 
to dryer filter cakes or shorter cycles 
on chemically conditioned slurries ( 22 ) . 

Pressure filtration is a fairly simple 
process. Initially, the slurry is chemr- 
ically treated and then pumped into the 
press (fig. 15). The filtrate flow from 
a large press can be 10,000 to 15,000 
gal/h, so that this part of the cycle is 
often called the fast fill. During this 
phase, the cake chambers of the filter 
press collect the major amount of slurry 
solids. As the chambers become progres- 
sively filled with solids, the pressure 
inside the press rises and the filtrate 
flow rapidly decreases. This portion 
of the cycle is the cake consolidation 



Filtrate 
outlet 



Feed Inlet 




Filtrate 
channels 



Filter 
medium 



Filter cake 



Recessed plates 



FIGURE 15, - Cross section of recessed-plate 
filter press (9, 22). 



phase, where slurry solids are being 
forced under pressure into the cake cham- 
bers. This compacting action displaces 
more water from the loosely formed slurry 
cake and enables the press to produce 
harder and drier cakes than other means 
of dewatering (6^, 22, 24). 

At a predetermined low-flow condition, 
pumping is stopped and the feed holes 
that carried the unconsolidated sludge 
into the press are blown clear. The 
stack of filter plates is undamped, and 
the individual plates are mechanically 
separated, allowing the filter cake be- 
tween adjacent plates to drop out of the 
chambers. The cakes then fall against 
breaker bars that reduce them to a con- 
venient size before they are transported 
for final disposal (6^, 9^). 

Although pressure filtration is a batch 
process, a series of presses can be or- 
ganized to work in a semi continuous oper- 
ation. This is done by having one press 
being filled, while a second is being 
airblown and a third is being unloaded or 
waiting to be filled again ( 25 ) . 

Filter presses find wide application in 
the process industries for the separation 
of slow-settling solids from liquids when 
the solids content is 1 to 10 pet and the 
f ilterability is poor. The use of these 
presses depends mainly on the particle 
size and on the quantity of solids in the 
slurry feed (22, p. 211). The capacity 
of the press is dependent on such factors 
as the characteristics of the material 
being filtered, operating temperature, 
pressure, filter medium being used, and 
type and size of filter press. There is 
no mathematical formula or scientific 
method for determining the exact filtra- 
tion capabilities or rates. The most re- 
liable method is to use past experience 
or records and to make controlled tests 
on laboratory-sized equipment (29). 

ADVANTAGES AND DISADVANTAGES OF 
FILTER PRESSES 

Filter presses are efficient at de- 
watering very fine grained solids such as 
slimes and consume only moderate amounts 



24 



of power during their operation. Labor 
requirements are high, however, and the 
batch operation may be a limitation In 
some facilities. Chemical pre treatment 
Is frequently needed to obtain the neces- 
sary solids capture. Also, large space 
requirements and the possible need to 
treat the decant liquors may restrict use 
(24) . Despite these problems , a well-de- 
slgned system of multiple filter presses 
might be the answer to difficult slime 
dewatering problems. 

CONTINUOUS PRESSURE FILTERS 

Although they are not well known, some 
pressure filters are available that work 
continuously. These units operate simi- 
larly to a drum or disk vacuum filter, 
except that the entire mechanism is en- 
closed in a pressurized tank. The cake 
is removed from the filter medium by a 
blade under superatmospherlc pressure. 
The dislodged cake is throttled out of 
the tank to atmospheric pressure by a 
self-sealing screw conveyor or a series 
of receivers and pressure locks ( 23 , 
p. 19-75). 

The operating pressure in these ma- 
chines is normally 100 Ib/ln^ with a 
pressure drop of 40 Ib/in^ across the 
filter medium. Filter areas from 4 to 
700 ft^ are available. The cost of con- 
tinuous pressure filters can be two to 
four times that of a vacuum filtration 
unit having the same filter area ( 23 , 
p. 19-75). 

Continuous pressure filters are advan- 
tageous because they work continuously 
and do not require close operator super- 
vision. These units are also capable of 
much higher filtration rates for low- 
compressibility solids than are vacuum 
filters. There are drawbacks, too, be- 
cause these machines are mechanically 
complex, do not allow access for any 
maintenance during operation, and have 
problems with lubrication. As mentioned 
before, these units are more expensive 
than similarly sized vacuum filters (23, 
p. 19-75). 



BELT FILTER PRESSES 

Belt filter presses have been manufac- 
tured in Europe since the mldsixtles and 
have been used mostly in pulp or paper 
plants and for sewage sludge dewatering. 
Only in the last several years have these 
machines been used by the minerals indus- 
try. Experimental work with the belt 
press indicates that this machine pro- 
duces a drier cake with lower polymer 
consumption than centrifuges or similar 
dewatering devices. The belt press cake 
also has a relatively high shear 
strength, which is important if the cake 
Is to be transported to a waste dump 
(28). 

The belt filtration process is composed 
of three operational stages: chemical 
conditioning of the feed slurry, gravity 
drainage to a nonfluid consistency, and 
compaction of the dewatered solids. Fig- 
ure 16 depicts a simple belt press and 
shows the location of the three stages. 
Although present-day belt presses may 
be very complex pieces of equipment, they 
follow the same concepts indicated in 
this figure. Good chemical conditioning 
is of primary Importance to successful 
and consistent dewatering by the belt 
filter press. The recent developments 
in high-grade polymeric reagents have 
encouraged the development of belt press- 
es to their present performance levels 
(19). 



Upper belt wash 



Chemically 
treated feed 




Drive rollers 



Dry cake 



FIGURE 16, - Cross section of belt filter press (28), 



25 



The first stage in belt press dewater- 
ing is to add a flocculating agent to the 
slurry in a conditioning zone (12). The 
resultant mixture must be allowed to set 
under low shear conditions to allow floc- 
culation to take place. If this mixture 
is disturbed during f locculation, more 
chemical agents may be needed to obtain 
the necessary degree of flocculation and 
any economic advantage will be lost. The 
conditioning zone should have an adjusta- 
ble residence time which can be increased 
or decreased to suit the slurry being 
treated. Correct conditioning liberates 
the free water originally contained in 
the slurry and also facilitates the cap- 
ture of the finest particles present by 
causing them to flocculate with the 
larger particles. This is necessary 
since water removal will be carried out 
on belts whose openings are larger than 
the smallest slurry particles (19). 

The next stage after chemical condi- 
tioning is gravity dewatering, which al- 
lows the free water to drain from the 
solid particles. The slope of the belt 
will normally be set at 5° to 15° from 
horizontal to assure proper distribution 
of the solids across the belt as the wa- 
ter is separated from the slurry. At the 
end of the drainage section, the slurry 
will lose about 50 pet of its free water. 
At this time, the formation of an even 
surface cake is essential for the suc- 
cessful completion of the following stage 
in the cycle. The even surface prevents 
uneven belt tension or distortion, and 
the rigidity of the semicake enhances 
distribution through the process (19) . 

Following the gravity dewatering stage, 
the semicake is sandwiched between the 
carrying belt below and the covering belt 
above. The semicake is squeezed between 
the two belts and subjected to flexing 
in opposite directions as it passes 
around various rollers. This action in- 
creases water release and allows greater 
compaction of the cake (19). At least 
one manufacturer has developed a high- 
performance belt press that has a low and 
a high belt pressure stage. This machine 



is basically the same as other belt 
presses except that in the high-pressure 
stage, an additional belt is added to in- 
crease the pressure on the solids cake 
and remove more water ( 12 ) . 

Finally, the dried cake passes from be- 
tween the two belts and is carried away 
from the machine by truck or conveyor 
belt for ultimate disposal (19). 

The material and construction of the 
filter belts themselves are very impor- 
tant. Belt filter presses do not utilize 
the belt fabric as a filter medium but as 
a support for the self -filtering solid- 
liquid mass. The mesh size of the belt 
fabric is not critical because very large 
floes will be produced by the polyelec- 
trolytes in the first stage of operation. 
Texture, however, is a significant fea- 
ture of the belts because it influences 
the adhesion of the cake to the belt and 
the physical strength of the belt. Cur- 
rently, the available belts are those 
used in the papermaking industry, al- 
though a wider variety of special-purpose 
woven belts are becoming available. Most 
belts are woven from polyester monofila- 
ment, and many weaving patterns are 
available (19), 

A major problem with the belt press is 
that an uneven layer of sludge causes the 
belt to deviate from the desired tracking 
line. This occurs because the filter 
belts usually have short lengths and 
widths. Direct side pressure should not 
be used because guide rollers would cause 
serious wear on the edges of the belt. 
One system that has been devised uses a 
delicately balanced sensor flap which 
rests lightly on the edge of the belt. 
This flap is attached to the spindle of a 
rotary directional fluid-flow control 
valve. Displacement of this valve causes 
hydraulic fluid to flow to either one or 
the other end of a cylinder which sup- 
ports one end of the belt roller. The 
cylinder's movement causes the roller to 
slew across the line of travel of the 
belt, and belt tracking is induced rather 
than forced (19), 



26 



The openings in the belt are easily 
plugged by the solids particles so the 
belt must be washed continuously. The 
cont£imination of these perforations de- 
pends on the belt tension, pressure on 
the cake, complexity of belt weave, type 
of sludge being treated, and quantity and 
type of polyelectrolyte being used. 
Washing is done using water; however, the 
volume and pressure used vary from manu- 
facturer to manufacturer. The current 
trend is to recycle the filtrate as belt 
washing water. This is satisfactory if 
the spray system is well designed to al- 
low cleaning in the event of a blockage. 
At low solids concentrations the volume 
of filtrate produced will be adequate for 
belt washing. 

ADVANTAGES AND DISADVANTAGES 
OF BELT FILTER PRESSES 

Belt filter presses have several advan- 
tages, which include — 

1. Continuous operation. 

2. Low power consumption, as low as 
3.7 kW according to one manufacturer 
(12). 

3. Low consumption of chemical addi- 
tives; one manufacturer claims his unit 
uses only 25 to 50 pet that used by cen- 
trifuges for some applications (12). 

The disadvantages of belt filter press- 
es include — 

1. Problems with proper belt tracking 
(19). 

2. Shortened belt life if very coarse 
solids are dewatered. 

3. Difficulty in dewatering slurries 
having a very low solids content ( 28 ) . 

VACUUM FILTRATION EQUIPMENT 

Vacuum filtration equipment has been in 
use in this country for over 100 yr, and 
continuous vacuum machines represent the 
most prevalent type of filtration equip- 
ment seen in the mineral industry today. 
Their simple operation and need for a 



minimum of auxiliary equipment has made 
them popular for solid-liquid separation 
in a variety of industrial applications. 

Continuous vacuum filters separate the 
solids and liquid in a slurry by placing 
the filter in the slurry and applying 
suction behind the filter so that the 
water and solids are drawn to it. The 
solids are collected on the filtration 
surface, while the water is drawn through 
the filter and separated from the solids. 
The particles accumulate to form a cake 
on the filtering surface that is gradual- 
ly lifted from the slurry. The cake is 
removed from the filter by various means, 
and the filter is returned to the slurry 
to repeat the process ( 18 , p. 12-51). 

Vacuum filters must have several items 
of auxiliary equipment in order to func- 
tion properly. An agitator is needed to 
keep the slurry solids in suspension un- 
til they are drawn to the filter medium 
( 32 , p. 226). A vacuum pump provides the 
subatmospheric pressure to draw the liq- 
uid through the filter. A receiver with 
a barometric leg is located between the 
filter and the vacuum pump and separates 
the liquid from air drawn through the fi- 
lter. Last, a filtrate pump removes 
the liquid from the receiver and forces 
it to the disposal point or back into the 
circuit for reuse ( 19 ) . 

The following sections describe the 
filtration equipment found in the min- 
erals industry. 

DRUM VACUUM FILTERS 

Drum vacuum filters are probably the 
most common type of continuous vacuum 
filters in use. They consist of a hori- 
zontal cylinder or drum whose circumfer- 
ence holds the filter medium. The drum 
rotates in a slurry tank, and a vacuum 
inside the drum draws the liquid through 
the unit, leaving the solids captured on 
the filter surface. Scrapers or other 
devices then remove the cake from the 
drum, and the medium reenters the tank, 
enabling operation to be continuous ( 24 ) . 
Figure 17 shows the front and side views 
of a typical drum vacuum filter. 



27 



SIDE VIEW 



FRONT VIEW 



Flltsr cake 




FIGURE 17» - Front and side views of drum vacuum 
filter (20). 



When a precoat method is used, an 
advancing-knife mechanism must be pre- 
cisely positioned to dislodge the cake 
and a minimum amount of the precoat ma- 
terial from the drum surface. This is an 
exacting procedure because the knife- 
advance system must not move more than 
several thousandths of an inch. The 
knife system has to be constructed so 
that its linear integrity, with respect 
to the drum face, is absolute and any de- 
viations from the correct distance must 
be extremely small (20, p. 1477). 



Vacuum drums can handle throughputs up 
to 250 gal/min, particles as small as 
0.0020 in, and solids loadings between 8 
and 10 pet. The average yield is between 
2 and 10 lb of dry solids per square foot 
per hour (24) . 

Drum vacuum filters have been widely 
used in industry because of their adapta- 
bility with respect to filter media and 
cake discharge methods. A variety of 
filter media can be easily used with a 
drum filter because the material itself 
is attached around the circumference of 
the cylinder. This configuration simpli- 
fies installation, inspection, mainte- 
nance, and removal of the filter medium. 
(24). 

CAKE DISCHARGE METHODS 

In addition to the variety of filter 
media available, there are many methods 
of removing the cake from the filter sur- 
face of a drum vacuum filter. The 
scraper-blade method of discharge is com- 
monly used where the cake is friable and 
has poor mechanical qualities. A cake 
that is fairly thick and is not strongly 
held to the filter cloth could be easily 
dislodged by a blade. This blade could 
be steel, rubber coated, or made of 
other materials compatible with the par- 
ticular corrosive or abrasive qualities 
of the filtrate or the solids. In many 
cases, a pressure reversal or blow is 
sufficient to loosen the cake from the 
filter medium, and the blade itself only 
guides the loose cake to the discharge 
chute (20, p. 1476). 



A roller discharge device may be used 
where a thin cake of sticky or thixotro- 
pic material is formed on the drum. This 
system uses a roller positioned close to 
the drum filter at the point of cake dis- 
charge. The cake is transferred from the 
filter to the roller and then removed 
from the roller by a cutting blade ( 20 , 
p. 1476). 

The belt discharge arrangement uses 
a filter cloth that winds around the 
drum to form the cake but leaves the 
drum to carry the solids to the dis- 
charged area. The cloth or belt passes 
over a small-diameter roller which causes 
the cake to separate from the filter. 
This system can be used for a cake hav- 
ing mechanical strength, for a thin 
cake, or where intensive washing of the 
filter is necessary to maintain the 
cloth openings. After the belt has 
passed over the discharge roll, it re- 
turns to the drum through a belt-washing 
system (20, p. 1476). Some of the vari- 
ous discharge methods are illustrated in 
figure 18. 

EQUIPMENT MODIFICATIONS 

Drum vacuum filters have been modified 
to separate solids and liquids under ad- 
verse conditions. For example, these ma- 
chines can be fitted with simple hoods 
which limit the escape of poisonous or 
foul-smelling vapors. They can be adapt- 
ed for complete sealing and for operation 
in a nitrogen environment; however, this 
complicates access to the internal parts , 



28 




Compressed 
air line 



Vacuum lines KNIFE DISCHARGE 



Atmospheric 
port 

Vacuum lines ROLLER DISCHARGE 



Atmospheric 
port 

Vacuum lines BELT DISCHARGE 



Vacuum lines 



PRECOAT 



Slurry 



FIGURE 18, - Schematic drawings of various dis- 
charge methods for drum vacuum filters (20), 

and the drum design must be such that the 
vacuum system and the cake receiving sys- 
tem are isolated to prevent vapor loss 
(20, p. 1481). 

The drum vacuum filter can also be mod- 
ified with accessories that improve the 
quality of the cake with respect to its 
washing and drying characteristics. This 
is possible because the cake moves 
through the washing and drying zones in 
the form of a continuous sheet and be- 
cause the cake and filter medium are ade- 
quately supported on the drum shell. The 
filters can be fitted with simple rollers 
which extend the full width of the filter 
drum and can be arranged to eliminate 
irregularities or cracks in the cake 
prior to washing and drying. The wash 
waters and air are therefore applied to a 
uniform surface and will not short cir- 
cuit or channel the deposited solids. 
The cake compression system may also have 
a "wash blanket" draped over the cake, 
which further limits any tendency for the 
air or wash water to channel the solids. 
These blankets allow the wash water to be 
used near the point where the cake 
emerges from the slurry ( 20 , p. 1481). 

Drum vacuum filters can also be de- 
signed so that the trough and hood are 
thermally insulated. If needed, the unit 



can be equipped with heat exchange equip- 
ment for heating or cooling the cake ( 20 , 
p. 1481). 

The basic design of the drum vacuum 
filter has undergone little change; how- 
ever, with the development of new and 
improved construction materials, many 
options have become available. Drums 
can now be fabricated in a variety of 
metals, plastics, and rubber for hand- 
ling corrosive materials. Internal pip- 
ing and valves have also been improved 
by the use of corrosion- and abrasion- 
resistant materials, which greatly reduce 
shutdowns for repair of leaks or loss of 
vacuum. Tanks and agitators can also be 
fabricated of special materials, and tank 
linings have greatly reduced wear (20 , 
p. 1475). 

PERFORMANCE 

For any given feed condition, the per- 
formance of drum vacuum filters can be 
optimized by the drum speed, the vacuum, 
and the percentage of drum surface sub- 
merged in the feed slurry. Most drum 
filters have controls for the manual ad- 
justment of these variables, and some 
models have automatic adjustments which 
are actuated by changes in the quality 
and quantity of the feed or cake ( 20 , 
p. 1480). 

A drum filter should be operated with 
the greatest degree of submergence and at 
the highest drum speed in order to maxi- 
mize the throughput of solids. It should 
be remembered that any increase in the 
submergence reduces the proportion of the 
drum area available for washing and dry- 
ing. Drum submergence above 40 pet ne- 
cessitates the use of seals on the drum 
shaft where it passes through the trough. 
No matter what combinations of drum 
speed, vacuum, and submergence are used, 
the sum effect must produce a cake that 
can be completely and easily removed from 
the drum. If this cannot be achieved, 
the filter medium will quickly deterio- 
rate and its life will be severely short- 
ened (20, pp. 1480-1481). 



29 



ADVANTAGES AND DISADVANTAGES OF 
DRUM VACUUM FILTERS 

The advantages of the rotary drum vac- 
uiua filter are — 

1. Continuous operation, which results 
in low operating labor costs. 

2. Many design and operational varia- 
tions available for a wide range of sus- 
pensions of divergent nature. 

3. Clean operation. 

4. Low maintenance costs. 

5. Effective washing and dewatering 
(32, p. 226). 

6. Provides a filtrate with a low sus- 
pended solids concentration. 

7. Does not require skilled personnel, 

8. Has low maintenance requirements 
for continuously operating equipment 
(19). 

The disadvantages of the rotary drum 
vacuum filter are — 

1. High capital cost. 

2. Limitations imposed by vapor pres- 
sure of hot or volatile liquids. 

3. Incapable of handling products that 
form explosive or inflammable gases under 
vacuum. 



4. Unsuitable 
slurries. 



for quick-settling 



5. Tendency for cloth blinding due to 
thin cakes and short cycles , although 
this may be alleviated by the applica- 
tion of a belt or string discharge ( 32 , 
p. 227). 

6. Auxiliary equipment such as vacuum 
pumps is very loud. 



7. Consumes the largest amount of en- 
ergy per unit of slurry dewatered, in 
most applications. 

8. Produces wetter cakes if blowback 
is used, and greater filter medium wear 
if blowback is used in conjunction with a 
scraper knife (32^, p. 227). 

ROTARY DISK VACUUM FILTERS 

The principle of operation for the 
rotary disk vacuum filter is the same as 
that for the rotary drum filter. The 
disks are oriented in a vertical plane 
and are composed of several pie-shaped 
sectors which fit into a central pipe for 
support of the disks and for transport of 
the filtrate. These sectors can be re- 
moved without disturbing the others in 
the same disk, and at slow speeds, it is 
possible to change a sector while the 
filter is still in operation. The filter 
medium cloth can be slipped over each 
sector and fastened at the innermost end 
of each sector ( 20 ) . As the disks ro- 
tate, they go through pickup and dewater- 
ing operations similar to those carried 
out on the drum filters. At the dis- 
charge point , the cake is removed by 
means of wires or knives. Figure 19 
shows two simplified views of a disk 
filter. Disk vacuum filters are avail- 
able with areas from 0.5 ft^ to approxi- 
mately 3,300 ft^; for large areas, as 
many as 12 disks are used in a single 
unit (32, p. 229). 

Owing to the vertical disk orientation, 
the wide variety of discharge methods 
used on drum filters cannot be applied to 
disk filters. Thus, their application is 
somewhat more limited than that of the 
drum filters, but a disk filter will oc- 
cupy only one-third the floorspace of a 
drum filter having the same filter medium 
area and is less expensive ( 20 , p. 1477). 

Disk filter provide more efficient agi- 
tation than drum filter agitators. In 
some applications a disk filter may be 
more cost efficient (18, p. 12-63). 



30 



SIDE VIEW 



FRONT VIEW 




Filter cakes 



Compressed 
air port 

Vacuum 



Scraper blades 

Slurry level 

Filter surface 

Filter sector 




FIGURE 19, - Front and side views of disk vacuum 
filter (13). 



ADVANTAGES AND DISADVANTAGES OF 
ROTARY DISK VACUUM FILTERS 

The advantages of using a rotary disk 
vacuum filter are — 



the liquid in the slurry feed passes 
through a filter, leaving the solids de- 
posited on top of the media. These units 
can be characterized as flat-bed, high- 
capacity filters which lend themselves to 
granular, fast-filtering materials and 
high-specific-gravity concentrates ( 20 , 
p. 1478). 

ROTARY TABLE VACUUM FILTERS 

The rotary table units have a circular 
shape, the filter medium and supports ro- 
tate about a central axis. The feed 
slurry is deposited along the radius of 
the unit and rotates while it is being 
subjected to vacuum dewatering, washing, 
and drying, finally being removed from 
the filter medium by mechanical means 
(20, p, 1478) , Figure 20 shows a cross 
section of a typical horizontal table 
filter. 



1, Low capital costs per unit area. 

2, Large filter areas with minimum 
f loorspace. 

3, Rapid medivim replacement ( 32 , 
p. 229). 

The disadvantages of the rotary disk 
vacuum filtration system are — 

1, Difficulties in washing the cake. 

2, Difficulties in discharging very 
thin cakes, 

3, Inflexible operation, 

4, High rate of medium wear with a 
scraper discharge, 

5, Unsuitability for noncoherent cakes 
(32, p, 229). 

HORIZONTAL CONTINUOUS VACUUM FILTERS 

These vacuum filters use a horizontal 
filter surface in the form of a table, a 
belt, or multiple pans in a circular ar- 
rangement. The operating principle of 
the horizontal filter is the same as that 
of the rotary drum or disk filter, where 



The rotary table machines permit a 
choice of cake thickness, washing time, 
and drying cycle where 4- to 5-in-thick 
cakes can often be handled. Sharp sep- 
arations between countercurrent wash 
waters are also possible because of the 
horizontal drainage configuration. Ro- 
tary table filters, though used in indus- 
try, are better suited for dewatering 
free-draining solids and not sticky or 
thixotropic slimes because of the diffi- 
culties in removing the solids from the 
filter medium (20, p. 1478). 



Discharge 
zone 



TOP VIEW 



SIDE VIEW 




Wash bar 
Filter cake 



Filler mediun 
and table 



FIGURE 20. - Plan and cross-section views of hor- 
izontal rotary vacuum filter (20). 



31 



HORIZONTAL BELT VACUUM FILTERS 

The horizontal belt vacuum filter uses 
an endless belt of filter fabric support- 
ed by a slotted or perforated endless 
belt. Both belts travel over one or more 
vacuum zones. The slurry is deposited 
onto the filter at one end, wash water is 
applied at one or more points along the 
path of belt travel, and the cake is 
dumped at the other end. The support 
belt and the filter are parted and di- 
rected along separate lines of pulleys. 
The filter is washed and rejoins the sup- 
port belt just ahead of the slurry depo- 
sition point. Figure 21 shows a cross 
section of a typical horizontal belt vac- 
uum filter ( 15 ; 20, p, 1478). 

The horizontal belt machines have a 
high capacity per square foot of area un- 
der vacuum, similar to the horizontal ro- 
taries. They are well adapted for a 
countercurrent discharge circuit and en- 
able the cake to be flooded with wash 
solvent so that it can be steeped in the 
wash liquid. They are suited for coun- 
tercurrent leaching or washing ( 20 , 
p, 1478). 

Horizontal belt filters have been manu- 
factured for more than 30 yr, but recent- 
ly have they been used for large- tonnage 
applications that require filter cake 
washing. For example, the largest unit 
available in 1950 had only 40 ft^ of 



active filter area, while modern machines 
have over 900 ft^ of filter area. Many 
mechanical improvements have also been 
made on these machines , which have im- 
proved their reliability. The biggest 
single improvement, though, was the abil- 
ity to manufacture a continuous drainage 
belt to very close tolerances , which 
spurred the development of large-capacity 
units (20^, p, 1479), 

The horizontal filter shows several ad- 
vantages because the slurry deposition on 
the horizontal filter belt eliminates the 
constant slurry agitation necessary for 
rotary types. In the horizontal filter, 
the cake travels as a ribbon of unwashed 
cake which is gradually washed to the re- 
quired purity, either concurrently or 
countercurrently (19) , 

Another advantage of the horizontal 
machine is its ability to resist filter 
medium blinding. Most slurries contain, 
in addition to some medium-sized mate- 
rial, a fraction of fines. No matter how 
efficient the agitation with the rotary 
drum, these fines will always be concen- 
trated near the surface within the 
trough. Consequently, the fine material 
is sucked against the filter cloth of the 
rotary drum filter and causes blinding 
before the large particles reach the 
cloth. This does not happen on hori- 
zontal belt filters because the biggest 
particles reach the filter cloth first. 



Slurry-] 



Wash 
dam 



, a a m y r- 



Wash water 



Wash dam 



^ 



Air box -^ 



Discharged liquid-* 



X 




Dewatered 
slurry 



*^o, 



scharged 
solids 



Filter cloth ^Carrier belt ^Filter 

cloth wash 

FIGURE 21. - Cross section of typical horizontal belt vacuum filter (]5, 20). 



32 



followed by the finer particles. As a 
result, the larger particles act as a 
precoat for the finer solids (19). 

Another concern with the rotary drum 
and belt filters involves maintaining a 
vacuum. Many cakes, after being de- 
watered, shrink and crack. This allows 
air to pass freely through the cake, 
which reduces the vacuum and the dewater- 
ing effect on the cake. On the rotary 
drum filter, this is remedied by using 
expensive blankets , rakes , or squeeze 
rollers. On the belt filter, a simple 
sheet of impervious material such as 
polyethylene trailed over the cake is 
usually adequate to maintain the maximum 
vacuum (19). 

ADVANTAGES AND DISADVANTAGES OF 
HORIZONTAL VACUUM FILTERS 

The advantages of the horizontal vacuum 
filters are — 

1. Excellent wash capability. 

2. Flexible operation. 

3. High-volume operation for fast-set- 
tling solids (22, p. 230). 

The disadvantages of the horizontal 
vacuum filters are — 

1. Requirements for large floorspace. 

2. High initial costs. 

3. Unsuitability for slow settling 
solids ( 32 , p. 230). 

SELECTING AND SIZING 
FILTRATION EQUIPMENT 

In the selection of filtration equip- 
ment, the job requirements must be com- 
pared to those associated with the equip- 
ment characteristics. Job-related fac- 
tors include slurry character, production 



magnitude, process conditions, performr- 
ance requirements, and permissible mate- 
rials of construction. The equipment- 
related factors are the type of cycle 
(batch or continuous) , driving force 
(gravity, pressure, or vacuum), produc- 
tion rates of largest and smarllest units, 
separation sharpness, washing capability, 
dependability, feasible materials of con- 
struction, and costs. This last item 
must include depreciation (installed cost 
plus expected equipment life) , mainte- 
nance, operating cost (labor, services, 
and filter media) , and penalty of product 
loss (if any). In addition, considera- 
tion must be given to preconditioning and 
the use of filter aids (20^, p. 1485). 
The suitability of the most common types 
of filters for various classes of slur- 
ries is summarized in table 1. 

Continuous filters are the most desira- 
ble when the process to which they con- 
tribute is a steady-level, continuous 
one; however, the rate at which the cake 
forms and the magnitude of production 
rate will probably be the critical fac- 
tors. For example, the use of a rotary 
vacuum filter is not practical if a 0.1- 
in-thick cake will not form under normal 
vacuum in less than 5 min and if more 
than 50 ft^/h of wet cake is to be pro- 
duced. The production use of batch fil- 
ters is harder to define, although they 
have been used in some processes that 
turn out 200 ton/d of dry solids. Occa- 
sionally, equipment flexibility and high 
filtering pressures will become more 
important than other factors that would 
otherwise dictate continuous equipment. 
Small-scale tests are essential for esti- 
mating the filtration rate, the washing 
characteristics, and other important fea- 
tures. Filtration is essentially an art 
rather than a science, and experience 
with the various aspects of vacuum fil- 
tration will help in better approaching 
the selection of equipment and evaluating 
test results (20, p. 1485). 



33 



TABLE 1. - Classification of selected vacuum filters (20, p. 1481) 
(Performance index: 1 = very poor or negligible; 9 = highest possible performance) 





Maximum 

area, 

ft2 




Slurry 




Relative performance 


Type 


classification' 


Cake 
dryness 


Cake 
washing 


Filtrate 




A 


B 


C 


D 


E 


clarity 


Horizontal belt^ ••••••. 


900 
160 

860 
860 
860 
860 
3,230 


X 
X 


X 
X 

X 
X 


X 
X 
X 

X 


X 
X 
X 


X 
X 


5-8 
4-7 

5-8 
5-6 
5-8 
NA 
2-3 


7-8 
8-9 

6 
5 
6 
6 

1 


6 


Horizontal rotary table^ 

Rotary drum: 

Knife discharge^ •••••••••••••• 


7 
8 


Roller discharge^ ....•...•.•.. 


8 


Belt discharge^ «•••••••••••••• 


7 


Precoat^ .••■•.••.•■...••...••. 


9 


Rotary disk^ 


6 



NA Not available. 

'a — High solids concentration (>20 pet), free draining, fast settling, high filtra- 
tion rates. 

B — Rapid cake formation, reasonably fast settling solids. 

C — Lower solids concentration, slow thin cake formation, difficult to discharge. 

D — Low solids concentration, slow cake formation, very poor strength properties. 

E — Very low solids concentration, solids usually blind normal filter media. 
2 For free-draining materials where good washing is necessary. 

^Wide range of types and sizes. Generally suitable for most slurries of types B 
and/or C. Can be fitted with various devices to improve washing and cake drying. 
^Suitable for slurries that blind most filter media. 
^Large throughput for small floorspace. 



LABORATORY TESTING 

Extensive research has been done to de- 
velop laboratory procedures for determin- 
ing the filterability of slurries. The 
filter-leaf test is the commonly used 



method of estimating the necessary filter 
area for a particular slurry. In this 
method, a test leaf is used which is cov- 
ered with a filter medixim identical to 
that intended for the full-scale filter. 
Figure 22 shows the typical apparatus 



Ring stand with platform to hold 
test leaf while drying 



Air bleed to reduce 
vacuum at leaf 



To vacuum 
source \ 



Sh 



utoff valve —^ i= 



Timer 



Rubber hose 




Stirring 
rod 



Gage reading in cubic 
feet per minute to 
measure air velocity 
during dry time 



Zl 



Filtrate flask, Iqt 
or larger 



Wash liquor ' — Slurry container, 

if required in 2-qt capacity 

graduated beaker 



FIGURE 22. - Typical laboratory installation for vacuum leaf tests (20). 



34 



needed for a leaf test, 
for the test follows: 



The procedure 



1, Condition 2 qt of sludge for fil- 
tration. The sludge should be thickened 
to the same concentration as the produc- 
tion slurry. 



operation. The test results will provide 
filtration parameters for the cake forma- 
tion, drying, and washing portions of the 
filtration cycle. The filter-leaf test 
is easy to perform; however, several pre- 
cautions should be observed to assure 
accurate results: 



2, Apply the desired vacuum to the 
filter leaf and immerse in the sample for 
1-1/2 min while maintaining sample agita- 
tion. The test leaf is usually inserted 
upside-down in the slurry to simulate the 
cake formation zone of a drum filter. 



1, Representative 
should be used. 



slurry samples 



2. The test should be repeated 5 to 10 
times to observe any filter medixim 
blinding. 



3, Bring the leaf to a vertical posi- 
tion and allow it to dry under vacuum for 
3 min. This simulates the cake draining 
and drying part of the cycle, 

4, Blow off the cake for 1-1/2 min, 
which gives a total drum cycle of 6 min. 
To discharge the cake, the leaf is 
disconnected from the vacuum, and air 
pressure of not more than 2 Ib/in^ is 
applied, 

5, Dry and weigh cake to determine 
percentage moisture, which can be com- 
puted from the equation — 



f _ Wds 

■Lev ~ 



-C V 



At I 



(18) 



where 



fcv = the filter cake formation 
rate, pounds per square 
foot per hour, 

W(js = the dry weight of the sol- 
ids cake, pounds, 

tcv = the cycle time, hours. 



3, The sample must be continually agi- 
tated to assure that it is homogeneous, 

4, The vacuum must be regulated so 
that it does not vary during the test. 
The vacuum should be the same as that in- 
tended for use in the full-scale opera- 
tion U, p, 11-29), 

After the tests have been completed, 
the results can be analyzed. A graph of 
the moisture content of the filter-leaf 
test cakes versus a correlating factor 
should be constructed. The correlating 
factor is calculated from: — 



where 



Fw = V 



a V 



td 



^(fe;) 



(19) 



Fv = the filter cake correlating 
factor, dimensionless , 



Vav = the 



volume of airflow 
through the cake per unit 
area of filtering surface, 
cubic feet per minute per 
square foot. 



and At I = the filter test leaf area, 
square feet (2^, p, 11-29), 

The test can be easily modified for other 
cycle times or discharge mechanisms. 
Filter leaves and testing instructions 
are available from most filter manufac- 
turers. It may be necessary to adjust 
the results obtained by a factor to 
coiiq)ensate for partial medium blinding 
and for scaling over a long period of 



and 



t(jv = the drying time, minutes, 

Pdc = the pressure differential, 
pounds per square inch, 
gage, 

Wdst = the weight of the dry 
cake solids for a given 
cake thickness , pounds per 
square foot (2^; JL8^, p. 11- 
31). 



35 



A decreasing moisture correlation indi- 
cates that the moisture content de- 
creases; and as the air rate through the 
cake per unit of filtering area is in- 
creased, the vacuum differential, or the 
length of the drying time, is increased. 
On the other hand, if the cake thickness 
and the cake weight are increased, the 
moisture content increases. Knowing the 
percentage of available drying time of 
the filter cycle and using the design 
formation, such as the proper cake thick- 
ness for a given type of filter, the vac- 
uum level, and the airflow rate through 
the cake, it is possible to predict for 
each cycle time the discharged filter 
cake moisture content expected from the 
full-scale filter. The filter area pro- 
vided in the design should be for the 
maximum solids removal rate plus a 5- to 
15-pct safety factor (2, p. 11-31). 

FACTORS AFFECTING FILTRATION 

Efficient vacuum filtration is influ- 
enced by many variables, of which some 
can be controlled and others cannot. The 
following items represent many of the 
factors that affect the final moisture 
content of a filter product: 

1. Cake thickness. 

2. Pressure drop across cake. 

3. Drying time. 

4. Volume of air or gas per minute 
per square foot of filtering area. 

5. Viscosity of filtrate. 

6. Surface tension of filtrate. 

7. Filter medium. 

8. Size distribution of solids, 

9. Permeability of cake. 

10. Specific gravity of dry solids. 



11. Inherent moisture of dry solids. 

12. Surface properties and other char- 
acteristics of solids. 

13. Type of filter and construction. 

14. Homogeneity of cake formation. 

15. Temperature of solids and gas ( 18 , 
p. 12-53). 

Two other conditions — feed solids concen- 
tration and cycle time — are important in 
vacuum filtration because they influence 
many of the above factors. 

The feed solids concentration is very 
important in the filtration process; con- 
sequently, a thickening device often pre- 
cedes the filter to ensure that the feed 
solids concentration is consistent with 
economic and efficient operation. A gen- 
eral plot of dry cake output versus feed 
solids concentration reveals a curve, as 
shown in figure 23 ( 18 , p. 12-54), 

Each slurry has its own characteristic 
filtration curve, which must be experi- 
mentally determined. In this example, 
the slurry exhibits a sharp incremental 
rate increase above 35 pet solids. Con- 
trolling the solids concentration be- 
tween the limits of 45 and 55 pet solids 



200 



< 
c 

z 
t- I 

< z 

O xi 



150 



100 — 



50 




10 20 30 40 50 60 

FEED SOLIDS CONCENTRATION, wt pet 

FIGURE 23, - Representative curve for cake forma- 
tion rate versus feed solids concentration (18), 



36 



will require less filtration area and 
the resulting filter operating costs will 
be reduced. Above 58 pet solids, this 
slurry becomes relatively viscous and 
transportation to the filter will be 
difficult. There is a point of inflec- 
tion at about 55 pet solids where the 
curve becomes asymptotic. This indicates 
that further slurry thickening above 55 
pet solids is impractical and uneconomi- 
cal because it produces only a slight 
increase in the filtration rate ( 18 , 
pp. 12-54 to 12-55). 

The other important consideration for 
filtration operations is the cycle time. 
In this discussion, cycle time will be 
concerned with rotary vacuum filters, al- 
though the same principles apply to other 
types of filters. 

The cycle time of a continuous vacuum 
filter is the amount of time the filter 
takes to make one complete revolution and 
is given in terms of minutes per revolu- 
tion. During each cycle, there are three 
phases of filter operation: cake forma- 
tion, cake dewatering, and cake dis- 
charge. At the end of each cycle, the 
filter discharges a certain amount of 
cake per given filter area. With these 
data, the dry cake formation rate can be 
expressed in pounds per hour per square 
foot of filtering area. A log-log plot 
of dry cake formation rate versus cycle 
time for an arbitrary feed solids concen- 
tration is shown in figure 24 ( 18 , p. 12- 
55). 

For this example, the slope of the 
curve is -0.5, based on an assumption 
that solids concentration and cake com- 
pressibility remain constant. This rela- 
tionship can be mathematically expressed 

as — 

t 



tcvo - the 



old 
minutes , 



cycle time. 



= f, 



cvn 



(20) 



where 



^cvo ~ the old cake formation 
rate, pounds per square 
foot per minute, 

fcvn - the new cake formation 
rate, pounds per square 
foot per minute. 



and 



tcvn ~ the new cycle time, min- 
utes (18, p. 12-55). 



Vacuum filters are normally equipped 
with variable-speed drives operating 
within a range of 1.5 to 9.0 min/rev. 
Thus, for any filter area, cake output 
can be doubled, tripled, or halved, as 
the situation requires ( 18 , p. 12-55). 

Cycle time also affects the filter cake 
moisture content and dischargeability. 
As a general rule, the filters should be 
sized for a cycle time of at least 3 min/ 
rev and preferably 4 min/rev. Cycle 
times less than 3 min/rev will increase 
the cake moisture and produce thin cakes, 
which are difficult to remove from the 
filter medium. Difficult cake discharge 
can mean sizable increases in filter 
maintenance costs ( 18 , pp. 12-55 to 12- 
56). 



200 



100 



< 

cc 50 

z 
o 

I- 
< 

(T 
O 

u. 

UJ 

< 
o 



10 







Slope -0.^^^^ 






^^ 



10 



CYCLE TIME, min/rev 



FIGURE 24, - Representative curve for cake forma- 
tion rote versus cycle time (18). 



37 



HYDROCYCLONES 



BACKGROUND 



Cyclones have found wide use in indus- 
try for solids classification and concen- 
tration. The cyclones designed for liq- 
uids are called hydrocyclones , hydraulic 
cyclones, or hydroclones. The basic sep- 
aration principle used in cyclones is 
centrifugal sedimentation, where the sus- 
pended particles are subjected to a cen- 
trifugal force which causes them to sep- 
arate from the fluid. Cyclones have no 
moving parts , and the necessary vortex 
motion is provided by the fluid itself 
(32, p, 101). 

Hydrocyclones , by themselves , are not 
capable of producing a dry solids product 
because they use liquids for their opera- 
tion. Cyclones are useful, however, in 
concentrating the solids content of a 
slurry ahead of another dewatering ma- 
chine, such as a filter or centrifuge. 

The cyclone consists of a short cylin- 
drical section attached to an inverted, 
truncated, conical section. The apex or 
bottom of the conical section is called 
the underflow orifice, A central over- 
flow orifice or vortex finder is fitted 
to the base of the cone, and a feed 
orifice is attached tangent ially to 
the cylindrical body section. Figure 25 
shows the cross section of a typical 
hydrocyclone. The slurry enters at high 



Overflow (liquid) 



Slurry feed 
Feed orifice 




Overflow orifice or 
vortex finder 



Conical body section 



Underflow orifice 



pressure through the tangential feed ori- 
fice into the cylindrical section, where 
a rotating force field is established. 
The solids in the slurry are settled to 
the side wall by this force, slide down 
the inclined wall to the apex of the 
cone, and exit through the underflow ori- 
fice. The liquid portion of the feed 
travels to the center of the cone with 
some of the finest solids and exits 
through the overflow orifice, A vortex 
of air extends throughout the length of 
the cyclone (^8, p, 12-26), 

A cyclone is a simply constructed de- 
vice; however, its principles of opera- 
tion are complex and there are many vari- 
ables that must be carefully evaluated to 
produce the desired separations. The 
most significant variables are — 

1, Cyclone diameter, 

2, Cyclone cone angle. 

3, Feed, overflow, and underflow ori- 
fice sizes. 

4, Length of cylindrical section, 

5, Feed pressure, 

6, Feed concentration. 



7, 
27), 



Particle size (18, pp. 12-26 to 12- 



■ Underflow (solids) 
FIGURE 25. - Cross section of hydrocyclone (1_8). 



The hydrocyclone diameter is the most 
important factor influencing the applica- 
tion and efficiency of a cyclone because 
the smaller the particle to be dewatered, 
the smaller the cyclone diameter that 
must be used. For most applications, the 
cyclone manufacturer will determine the 
size and cone angle of the cyclone to be 
used (^8, p. 12-27). 

The orifice size is another important 
factor that influences cyclone perform- 
ance. The underflow orifice will de- 
termine the concentration and flow of the 
thickened solids from the cyclone. 
Enlarging the underflow orifice Increases 
the flow rate and the percentage of fines 



38 



in the underflow. The larger underflow 
orifice also allows more liquid to pass 
through it and decreases the concen- 
tration of the product, A larger over- 
flow orifice increases the total flow, 
the concentration of solids , and the max- 
imum particle size contained in the over- 
flow of the cone, A secondary effect is 
that the underflow of the cone increases 
in solids concentration and contains a 
larger percentage of coarse sizes ( 18 , 
p, 12-27), 

Changes to the feed orifice affect the 
volume processed by the cone because as 
the area of the feed orifice is in- 
creased, there is an accompanying propor- 
tional increase in the flow to the cy- 
clone. The additional flow reduces the 
retention time of the slurry within the 
cone and causes the cyclone to reject 
coarser materials to the overflow. As a 
result, when a cyclone feed is increased, 
the cyclone underflow orifice should also 
be increased to accommodate the higher 
tonnage of material fed to the cyclone 
(JJ^, p, 12-28), 

The feed inlet can be either rectangu- 
lar or circular; however, a rectangular 
inlet with its long side parallel with 
the axis of the cyclone produces better 
results. The top of the feed orifice 
should be flush with the top of the cy- 
clone to eliminate a dead space which 
would cause short circuiting of the feed 
(32^, p. 113). 

In the mineral industry it is common 
practice to have several vortex finders 
with different diameters or nozzles which 
can be put into the exit pipe. This ena- 
bles the operator to change the length of 
the vortex finder, when needed. An in- 
crease in the length of the vortex finder 
improves the efficiency of removal of the 
coarse particles but decreases the effi- 
ciency for the finer particles ( 32 , 
p. 114), 

The feed pressure affects the volume 
processed and the relative efficiency of 
the cyclone. The total flow to a cyclone 
will vary proportionally to the square 
root of the pressure. Increasing the 



feed pressure causes the underflow con- 
centration to increase and become finer 
in size, while the overflow discharge in- 
creases and the overflow solids also be- 
come finer ( 18 ) , 

The interior surface of the cyclone 
should be as smooth as possible to pro- 
mote good material flow. Abrasion re- 
sistance should be built into a cyclone 
if it is to be operated with abrasive 
solids, A wide range of construction ma- 
terials, such as steel, nylon, ceramics, 
polyurethane , and rubber, are available 
(32, p, lU), 

ADVANTAGES AND DISADVANTAGES 
OF HYDROCYCLONES 

Hydrocyclones have been used for many 
solid-liquid separation applications be- 
cause of the following advantages: 

1, High capacity, 

2, Simple operation, 

3, Compact design, which uses a mini- 
mum of floorspace, 

4, Relatively low capital costs. 

5, Low maintenance and processing 
costs (8^), 

Hydrocyclones do have a drawback, how- 
ever, because the smallest particles in 
the slurry will be carried away with the 
overflow. Proper cyclone design can min- 
imize this loss, but clarification of 
this liquid may be necessary, and this 
possibility should be recognized by the 
mill operator, 

SIZING HYDROCYCLONES 

The designer of a hydrocyclone instal- 
lation will be concerned with the size 
and number of cyclones needed. This 
will be based on the desired separa- 
tion efficiency and flow rate. Several 
small-diameter cyclones working in paral- 
lel are more efficient than one large 
cyclone handling the same capacity ( 32 , 
p, 114), 



39 



A simple, straightforward method of 
sizing hydrocyclones is to first deter- 
mine the requirements of the installation 
in terms of particle size, amount of 
throughput, or pressure differential be- 
tween the inlet and overflow orifices 
( 32 , p. 114). Knowing this information, 
manufacturers' bulletins can be consult- 
ed. Operational data for hydrocyclones 
are often presented in graphic form; fig- 
ure 26 shows a plot of pressure drop ver- 
sus throughput capacity for two represen- 
tative hydrocyclones. As can be seen, 
these two cyclones are designed to handle 
very small solids. 



1 25 


_ 1 


1 Model 
/ 


-w 




1 


r" 


l_ 


1 00 





/ 













90 





A 






Model 


B 


— 


80 





r@ 






/ 




— 


70 


— 


/ 






/ 




— 


60 


— 


/ 






/ 




— 


50 


— 


L 






f@ 




— 


40 


1 


m 






/ 




- 


30 


- 






/ 






— 


25 


- 






/ 






— 


20 


J@ 




) 


^ 






- 


1 5 


— 




/ 








- 


1 


1 


1 %i 


1 

1 




1 




1 



KEY 

^, Particle size X 10"^ 
etc. inch for 95- pcf separation 
efficiency for a given 
pressure drop and 
tfiroughput capacity 



0.6 0.8 1.0 2 4 8 8 1 

THROUGHPUT C A P ACIT Y , ga I / tnln 



FIGURE 26.- Graph of pressure drop versus through- 
put capacity for two hydrocyclones (8), 



More sophisticated methods can be used 
to determine the diameter, cone angle, 
and other dimensions for a particular 
situation. Their application, however. 



is very complex and a reference such as 
Svarovsky (32, pp. 106-118) should be 
consulted for a complete theoretical 
discussion. 



THERMAL DEWATERING 



BACKGROUND 

The use of heat is another method of 
separating liquids from solids. Thermal 
processes can be used to dry a slurry or 
just the thickened solids; however, fuel 
consumption will increase proportionally 
with the increase in the moisture pres- 
ent. It is, therefore, cost effective to 
use thermal drying after a substantial 
amount of moisture has already been re- 
moved by other methods such as centrifu- 
gation or filtration. 

There are basically three methods of 
heat transfer for drying: 



by convection or direct contact between 
the wet solids and hot air. Over the 
years, many different configurations of 
convection dryers have been available for 
the minerals industry. Now, six basic 
types exist, as follows: 

1. Drum. 

2. Suspension or flash, 

3. Multi louvre. 

4. Vertical tray. 

5. Continuous carrier. 



1 . Convection — the direct contact of 
particles with warm air. 

2. Conduction — the direct contact of 
particles with a heated shell of a dryer 
or other heated particles. 

3. Radiation — heat radiating from a 
hot surface to the particles ( 13 , p. 27- 
75). 



6. Fluidized bed (^, p. 13-7). 

In each of these dryers, the wet solids 
are separated from each other and sub- 
jected to a flow of hot gases or air. In 
the drum-type dryer, the solids enter one 
end of a rotating cylinder through which 
the hot air is blowing. The solids are 
tumbled the length of the cylinder and 
exit the other end (13, p. 27-76). 



THERMAL DRYER OPERATION 

Dryers used commercially for drying 
minerals commonly utilize heat transfer 



In the suspension or flash dryer, the 
wet solids are carried upward through a 
vertical duct by a blast of superheated 
air for a very brief time. The air is 



40 



usually heated to 1,200° F, but the sol- 
ids are in contact with the hot air for 
only a fraction of a second and are not 
changed chemically ( 18 , p. 13-39). 

The louvered arrangement has a series 
of specially designed flights which carry 
the solids upward through a flow of hot 
air. The vertical tray, on the other 
hand, has a series of shelves placed in a 
terrace arrangement. The solids are fed 
in at the top, and a vibrating action 
causes them to fall from one shelf to the 
next through heated air. The continuous 
carrier type uses a vibrating inclined 
screen to support the solids as they tum- 
ble through the hot airstream ( 18 , 
pp. 13-39 to 13-43). 

The f luidized-bed dryer uses a slightly 
different approach and has seen wider use 
in the past two decades. In this type of 
dryer, the solids are fed into a heater 
box and subjected to a high-velocity flow 
of hot air. The violent bubbling action 
of the solids is similar to that of boil- 
ing water, leading to the description of 
the dryer as a f luidized bed (4^) . Fig- 
ures 27 and 28 show simplified cross sec- 
tions of the various thermal dryers. 



Rotating drum 



Dry product 




DRUM 



Vertical 

draft tube SUSPENSION 



Falling material 
Dry produc 



MULTILOUVER 



Wet feed 



FIGURE 27, - Simplified cross sections showing 
operation of drum, suspension, and multilouver ther- 
mal dryers (18), 



Wet feed 



Vibrating tray 




Hot air 



Wet feed 



Vibrating screen 




VERTICAL TRAY 



Dry product 



CONTINUOUS CARRIER 



Dry product 



ADVANTAGES AND DISADVANTAGES OF 
THERMAL DRYERS 

Thermal dewatering equipment has sev- 
eral advantages , which include — 

1. Ability to reduce the moisture con- 
tent of slimes to 6 pet or less. 



Wet feed ►>:..; 



FLUIDIZED BED 



T ^Perforated plate 



2. Minimal labor requirements. 

3. Capability to operate continuously 
as long as feed material is available, 

4. Low maintenance costs. 



FIGURE 28, - Simplified cross sections showing 
operation of vertical tray, continuous carrier, and 
fluidized-bed thermal dryers (18), 

too low, the powdery product may be dif- 
ficult to handle or transport (1). 



There are, however, several disadvan- 
tages, such as — 



2, The fuel demands can become very 
high (4). 



1. The moisture content of the product 
must be carefully monitored. If it is 



3. A mechanical dewatering device is 
usually needed ahead of a thermal dryer. 



41 



SIZING THERMAL DRYERS 

As wet solids are dried, they go 
through three stages with respect to 
moisture loss (13, p. 27-75) : 

1. Warming period. 

2. Constant-rate period. 

3. Falling-rate period. 

The constant-rate and falling-rate 
periods have the greatest impact on the 
drying time required for a particular 
material. The warming period will be 
significant if radiant heat from sur- 
rounding surfaces is negligible ( 13 , 
p. 27-75). The rate of liquid evapora- 
tion during the constant rate period can 
be estimated using mass transfer or heat 
transfer equations as follows: 

Ec = Ma Ky (Hai - Ha) A^j 

(mass transfer), (21) 

17 _ hy (Ta - Ti) Ad 
^'-- ^^ 

(heat transfer) , (22) 

where Ec = the rate of evaporation, 
pounds per hour. 

Ma = the molecular weight of air, 
pounds per pound-mole. 

Ky = a mass transfer coefficient, 
pound-moles per square foot 
per hour. 

Ha = the humidity of the ambient 
air, pounds of water per 
pound of dry air. 

Ha I = the humidity of the air at 
the solid-air interface, 
pounds of water per pound 
of dry air. 



A(j = the drying area, 
feet. 



square 



and 



h„ = a 



I heat transfer coefficient, 
Btu per square foot per 
hour per °F. 



For airflow parallel with surface of 
solid, hy = (0.0128 Vam)°*®. 

For airflow perpendicular to surface of 
solid, hy = (0.37 Vam)°-^^. 

Vam = the A±r mass velocity, pounds 
per square foot per hour. 



Tg = the temperature of the ambi- 
ent air, "F, 

Ti = the temperature of the air at 
the solid-air interface, "F, 



and X| = the latent heat of water at 
the temperature , T | , of the 
solid-air interface, Btu per 
pound (22, p. 27-75). 

Equations 21 and 22 are valid while the 
surface of each particle is saturated. 
When this condition is no longer the 
case, thermal drying enters the falling 
rate period (_33, p. 17-02). 

If the total average moisture content 
(Xfav) of ^^y particle is composed of the 
average free moisture (Xfav) ^^d chem- 
ically bound or equilibrium moisture 
(Xeq), the rate of liquid evaporation can 
then be estimated by the liquid diffusion 
equation, which is — 



_ Xfav Xeq _ Xf 



Xti - X 



av 



eq 



Xfi 



= !^G""'^|"""'^-)' (23) 
where "^ ~ ( T ) » 



e = 



Dmt, 



_ "m«-n 



Xti = 



S2 



the rate of liquid evapora- 
tion during the falling rate 
period, pounds of water per 
pound of dry solid, 

the initial total mois- 
ture content of the solids , 
pounds of water per pound of 
dry solid. 



42 



■■tav ~ 



■■eq 



the average total moisture 
content during the falling 
rate period, pounds of water 
per pound of dry solid, 

the equilibrium moisture con- 
tent of the solids, pounds 
of water per pound of dry 
solid. 



decreased. To design for the percent 
open area, a factor called the discharge 
coefficient must be determined with the 
aid of a single-orifice test plate that 
has the same cross-sectional characteris- 
tics as the proposed full-scale dryer. 
Information obtained from the test plate 
can be used in the following equation to 
determine this coefficient: 



Xfi = the initial free moisture 
content at the beginning 
of the falling rate period, 
pounds of water per pound of 
dry solid. 



Ch = 



qa Ma P 



op 



4825 Ao Pdo Tk 



(24) 



where C^ = the discharge coefficient, 
dimensionless , 



Xfav = the average free moisture 
content during the falling 
rate period, pounds of water 
per pound of dry solid, 

e = Euler's number - 2,71828, 

Dm = the diffusivity value of 
moisture through a solid, 
square feet per hour, 

tn = any arbitrary time after the 
beginning of the falling 
rate period, hours. 



qa = the airflow through the ori- 
fice plate for the air tem- 
perature and pressure used 
during the test, cubic feet 
per minute. 

Ma = the molecular weight of air, 
29, 

Pop = the mean absolute pressure 
in the orifice, pounds per 
square inch , 

Aq = orifice area, square feet. 



and S = one-half the thickness of the 
layer of solids in the dry- 
er, feet (13, p, 27-75). 



P(jo = the pressure differential 
across the orifice, inches 
of water. 



These equations are helpful in estimat- 
ing the rate of evaporation of moisture 
and will enable the designer to determine 
the total drying time, feed rate, air- 
volume requirements, and other important 
factors concerning thermal dryers. 



When a 
considered 
orifice 
the plate 
most appli 
to 10 or 
Pressure 
than this 
ing gases 
p, 13-18). 



f luidized-bed dryer is being 
the pressure drop across the 
plate must be evaluated and 
size designed accordingly. For 
cations, a pressure drop equal 
12 in of water is adequate, 

differences slightly higher 
should be used if the fluidiz- 

are as hot as 1,200° F (18, 



The pressure drop will increase as 
the percent open area in the plate is 



and Tk = the temperature, 
(18, p, 13-17), 



kelvins 



After the discharge coefficient has 
been determined for a particular orifice 
configuration, the volume of air needed 
to pass through the orifice to provide 
the necessary pressure drop can be calcu- 
lated. The total number of orifices for 
the full-scale dryer can then be found if 
the designer knows the volume of air 
needed to maintain the pressure drop for 
one orifice and the total volume of air 
needed for drying the solids (18, p. 13- 
19). 

These equations will provide approxi- 
mate values for sizing thermal dryers, 
but as with other equipment , it is recom- 
mended that pilot-plant testing be done. 



43 



CURRENT BUREAU OF MINES RESEARCH ON DESLIMING METHODS 



The bulk of this review has been con- 
cerned with well established methods of 
dewatering slimes; however, at least two 
Bureau of Mines innovations for dewater- 
ing slimes should also be brought to 
the attention of the mining industry: 
Elect rokine tic methods and the rotary 
trommel. 

ELECTROKINETIC METHODS 

Background 

The Bureau is currently developing a 
method of dewatering slimes using an 
electrokinetic potential. This dewater- 
ing technique is intended to be used at 
the place of the slimes disposal rather 
than in the mill or other processing lo- 
cation. This method takes advantage of 
the electrical surface charge on the 
solid particles in a water suspension. 
Bureau of Mines research has shown this 
technique to be generally successful for 
treating siliceous mine tailings from 
several north Idaho metal mines , thicken- 
er underflow from two Appalachian coal 
preparation plants, and materials from 
numerous other coal and metal mine sites 
(30). 

Not all slimes or sludges can be de- 
watered by this method, however, because 
the sludge from an acid mine drainage 
treatment plant and scrubber sludge from 
a large coal-fired power plant were not 
responsive to this technique. The physi- 
cal properties of the slurries that af- 
fect their response to treatment include 
electrical conductivity, particle-size 
distribution, water content, and surface 
charge density. Chemical properties also 
influence behavior , but their importance 
is specific to each sample and difficult 
to characterize (30) . 

Application 

The electrokinetic phenomena of elec- 
trophoresis and electro-osmosis are prin- 
cipally responsible for the effects ob- 
served. Electrophoresis is the migration 
of small electrically charged particles 



through a stationary liquid due to an ex- 
ternal electrical potential. Electro- 
osmosis, on the other hand, is the mi- 
gration of liquid through a stationary 
porous solid as a result of an external 
electrical potential. 

Particle sedimentation can be acceler- 
ated by imposing a properly oriented 
electrical field on a slurry. For exam- 
ple, in a slurry composed of negatively 
charged solids, the anode or positive 
electrode is located at the bottom of the 
slurry, and the cathode or negative elec- 
trode is positioned at the surface 
(fig. 29). 

Practically speaking, the anode can be 
a section of abandoned steel track or 
wire mesh placed on the floor of the fill 
area prior to the slurry deposition. The 
cathode can be wire mesh positioned at 
the slurry surface and suspended from 
wood floats or cables attached to the 
roof of the mine opening. A relatively 
large direct voltage of 2 to 6 V/in^ of 
tailings surface area is applied to the 
electrodes, and the negatively charged 
solids begin migrating downward to the 
positively charged anode. The electro- 
phoretic migration will effectively ac- 
celerate the settling of these solids 
(30). 



INITIAL PARTICLE SEPARATION 



Elec trode 
polarity 



-Top electrode (cathode) suspended 
from a suitable buoyant float 

■ C lear water 




yar'ticle migratio 



^ Bottom 



- Slurry 



electrode ( anode) 



FINAL SOLIDS DEWATERING 



Electrode 

p olarity 



Top electrode (anode) 



[fit f 

Water migration 

1 \Mn iM 



-Consolidated solids 
-Sand drain 



Bottom electrode (cathode) 



FIGURE 29. - Two configurations for elec- 
trokinetically dewatering slimes (30). 



44 



When sedimentation is complete, further 
drainage can be stimulated by reversing 
the polarity so that the bottom electrode 
becomes the negatively charged cathode. 
The water Immediately adjacent to the 
negatively charged solid particles will 
contain excess dissolved positive ions 
that, in effect, give this water a posi- 
tive charge. In an electrical field es- 
tablished between two separated elec- 
trodes burled in a slurry, the solid 
particles will not move appreciably be- 
cause of their relatively dense packing, 
but the water will be carried toward the 
negatively charged cathode by the vis- 
cous drag of the migrating positive ions. 
This movement of water by direct current 
potential is electro-osmosis. As the wa- 
ter migrates to the cathode, the liquid 
can be removed by using slotted pipes or 
gravel drains. 

Electrophoretic flow is relatively in- 
dependent of pore size and is particu- 
larly attractive for dewatering dense 
slurries of fine particles where hydrau- 
lic flow of water through the sediment 
is negligible because of small pore size. 
Feasibility of electrokinetlc dewater- 
ing for a particular slurry is best 
determined by direct testing in the 
laboratory. The complex Interaction of 
factors affecting the efficiency of the 
process has prevented the use of physical 
properties alone as reliable predictors 
of performance. Change in application 
methods, such as current density, current 
reversal, electrode configuration, or 
settlement time, can also have unique and 
important effects on the response of a 
given slurry (30) . 

Current Research and Use 

Field tests conducted in two Idaho 
mines demonstrated that the electrokinet- 
lc process can effectively dewater and 
denslfy unclassified mill tailings or 
slimes for use as backfill, with moderate 



power consumption. One mine is preparing 
to use the process to dewater slimes un- 
derground as a regular operating proce- 
dure, and another mine plans to use it 
when space occupied by old slime deposits 
needs to be recovered for other purposes 
(30). 

TROMMEL SCREEN 

Description of the Method 
and Equipment 

The Bureau is also doing extensive re- 
search on the trommel screen for dewater- 
ing slurries. In this method, the slurry 
is first mixed with a flocculating agent 
to agglomerate the small particles into 
much larger masses. Next, the treated 
slurry passes over a hydrosleve to remove 
some of the liquid. The remaining wet 
solids then go into the upper end of a 
long inclined cylinder or trommel made of 
steel screen. The trommel rotates about 
its long axis and allows any remaining 
liquid to pass through the screen while 
the dewatered solids move through the 
cylinder and exit through the lower end 
(fig. 30) (27_, 37). 

The most important aspect of this meth- 
od is adding the proper flocculant to the 
slurry. It must cause the solids to form 
sufficiently large masses that will not 
pass through the screen with the liquid 
but will remain on top of the screen sur- 
face. At the same time, the flocculating 
agent must be potent enough so that a 
minimum amount of the chemical will pro- 
duce the required f locculation. Exten- 
sive testing has indicated that polyethy- 
lene oxide (PEO) is suitable for use with 
many different slurry types in the trom- 
mel screen. This flocculating agent is a 
water-soluble polymer having a nominal 
molecular weight of 8 million. This 
agent causes slurry solids to flocculate 
within minutes and works well with the 
hydrosleve and trommel screen (27, 37). 



45 



Chemically 
treated slurry 



Rotating trommel 




Liquid discharge 



Solid discharge 



FIGURE 30, - Diagram showing operation of 
rotary trommel (27), 



and 9°. The slurry moved over the hydro- 
sieve and through the trommel by gravity 
(27, 37). 

Test Results 

Test results using this method indicate 
that phosphatic and coal-clay slurries 
can be successfully dewatered. In one 
particular test, a coal-clay slurry had 
23.9 wt pet solids, of which over 70 pet 
were smaller than 325 mesh. A solution 
containing 0.125 wt pet PEG was added to 
the slurry at a dose of 0.78 lb PEG per 
ton of dry solids. The dewatered product 
had 60.1 wt pet solids (37). This im- 
pressive value is representative of the 
results obtained so far. 

Conclusions 



The hydrosieve used for this research 
was constructed of stainless steel and 
was 8 ft long. The first 4 ft had screen 
openings of 0.04 in, and the last 4 ft 
had openings of 0.02 in. The hydrosieve 
was inclined at an angle of 58° from 
horizontal (37) . 

The trommel screen was also composed of 
stainless steel but had 10-mesh openings. 
The trommel had a length of 36 in and a 
diameter of 6 in. The angle of inclina- 
tion from horizontal was between 3° and 



Research on the trommel screen is on- 
going, so no information is available on 
production costs per ton of dried solids. 
The equipment is fairly simple in design 
and would not represent a substantial 
capital investment; however, this method 
does require chemical pretreatment that 
would significantly affect the cost of 
dewatering slimes. Nonetheless, this 
method shows promise for mineral industry 
use, and research is continuing on the 
refinement of the trommel screen. 



DISCUSSION 



In recent years , economic pressures 
have caused milling operations to max- 
imize the separation of economic min- 
erals from the waste material. This 
has been done by grinding the ore to 
much smaller particle sizes. This fine 
grinding, while increasing the mineral 
extraction, has also posed a serious 
disposal problem for the waste materi- 
al. As practically all mineral bene- 
ficiation involves the use of water, 
the resulting solid-liquid mixture or 
slurry must be properly disposed of in 
accordance with the current environmental 
regulations. 



There is an abundance of dewatering 
equipment available that can separate the 
solids from the water with varying de- 
grees of efficiency. Physical separation 
methods , such as gravity thickeners , cen- 
trifuges , filters , thermal dryers , and 
cyclones , all reduce the water content or 
increase the solids content of the slur- 
ry. Although these items were discussed 
separately, they can be used in conjunc- 
tion with each other to produce a solids 
mass with an acceptably low moisture con- 
tent. Figure 31 shows the relative capa- 
bilities of common pieces of dewatering 
equipment in terms of the slurry particle 
size. 



46 



o 
a 

UJ 

< 
O 

CO 

o 

_l 

o 

CO 



LU 
QC 

ID 
I- 
CO 



100.0 
80.0 

60.0 

40.0 

30.0 
25.0 

20.0 
15.0 

10.0 

8.0 
6.0 

4.0 

3.0 
2.5 
2.0 



325 200 



TYLER MESH 
100 48 28 20 14 10 



8 



1.5 — 



1 \ r 



6 

T 



Gravity thickeners 



Hydrocyciones 



Drum vacuum filters 
Solid-bowl centrifuges 



Electrokinetic methods 
Disk vacuum filters 
Belt presses 
Filter presses 



Horizontal vacuum filters 



Drum vacuum filters 



Solid-bowl centrifuges 



Vibrating-basket centrifuges 



Positive-discharge basket centrifuges 



1.0 



Thermal dryers 



1 



1 



1 



0.001 0.002 



0.005 0.01 0.02 0.05 

PARTICLE SIZE, in 



0.1 



0.25 



FIGURE 31. - Chart showing generalized capabilities of commonly used dewatering equipment with 
respect to solids cake, moisture content, and particle size (6, 18, 25, 28, 30-31). 



The Bureau is continuously doing re- 
search that will benefit the mining in- 
dustry. Classical dewatering methods are 
being improved, and new methods are being 
devised. Electrokinetics and the rotary 
trommel are but two ongoing Bureau proj- 
ects that should help the mining industry 
reduce costs and increase efficiency. 

Despite the amount of knowledge we have 
concerning dewatering, its practice is 
still an art. Each mine generates waste 
material which, in one way or another, is 
different from the waste of any other 
mine. Variations in mineral content and 



physical properties must be evaluated 
carefully so that the right combination 
of equipment and methods will produce a 
sufficiently dry product. This paper was 
prepared in order to give the mill opera- 
tor or owner an overview of what dewater- 
ing equipment is available. An awareness 
of the different dewatering practices 
will enable such operators to evaluate 
alternatives that perhaps would not have 
been considered otherwise. This paper 
should also serve as a basis for further 
detailed research into the many ways of 
meeting the desliming challenge. 



47 



REFERENCES 



1. Anderson, J. C. Coal Waste Dis- 
posal To Eliminate Tailings Ponds. 
Min. Cong. J., v. 61, No. 7, July 1975, 
pp. 42-45. 

2. Baker, M. , Jr., Inc. (Beaver, 
PA). FGD Sludge Disposal Manual. Rept. 
FP-977 (prepared for the Electric Power 
Research Institute, Palo Alto, CA) , Jan, 
1979, 536 pp. 

3. Bird Machine Co., Inc. (South 
Walpole, MA). Bird Centrifugals: for 
Clean Coal and Refuse Dewatering. 1980, 
6 pp. 

4. Casili, J. T. Heat Drying Sludge 
From Ponds. Min. Cong. J., v. 61, No, 1, 
Jan. 1975, pp. 34-37. 

5. Centrifugal and Mechanical, Inc. 
(St. Louis, MO). CMI Model EBW. Undat- 
ed, 4 pp. 

6. Coal Age. Improved Equipment 
Available Now. V. 85, No. 1, Jan. 1980, 
pp. 56-61. 

7. Cook, R, L. , and J. J. Childress. 
Performance of Lamella Thickeners in Coal 
Preparation Plants. Min. Eng. , v. 30, 
No. 5, May 1978, pp. 566-571. 

8. Dorr-Oliver, Inc. (Stamford, CT). 
Cyclones. 1976, 28 pp. 

9. Duriron Co., Inc., Filtration Sys- 
tems Division (Angola, NY). Durco Quadra 
Press Filters. Bull. EF/21, 1980, 7 pp. 

10. Emmett, R. C, and R. P. Klepper. 
Technology and Performance of the Hi- 
Capacity Thickeners. Min. Eng., v. 32, 
No. 8, Aug. 1980, pp. 1264-1269. 

11. Envirotech Corp. (Salt Lake City, 
UT), Sludge Dewatering for FGD Products. 
Rept. FP-937 (prepared for the Electric 
Power Research Institute, Palo Alto, CA, 
Apr. 1979, 260 pp. 

12. Fischer, M. C, and M. G. Schill. 
The Dewatering of Fine Coal Refuse With a 



Continuous High Performance Belt Filter 
Press. Pres, at Fall Meeting, Soc. Min. 
Eng,, AIME, Tucson, AZ, Oct. 18, 1979, 11 
pp.; available from the authors, Parkson 
Corp., Fort Lauderdale, FL. 

13. Given, I. A. (ed,). SME Mining 
Engineering Handbook. Society of Mining 
Engineers of AIME, 1973, 2666 pp. 

14. Jacobsen, S. P., W. Roushey, and 
E. L. Rau. Coal Waste Dewatering Systems 
(contract J0205012, CO Sch, Mines Res, 
Inst.). BuMines OFR 114-81, 1981, 133 
pp.; NTIS PB 81-244501. 

15. Joy Manufacturing Co. , Denver 
Equipment Division (Colorado Springs, 
CO) . Denver Horizontal Belt Vacuum Fil- 
ter. Bull. F 18-B103, 1979, 11 pp. 

16. Kealy, C. D. , R. A. Busch, and 
M. M. McDonald. Seepage-Environmental 
Analysis of the Slime Zone of a Tailings 
Pond, BuMines RI 7939, 1974, 89 pp. 

17. Keane, J. M. Sedimentation: The- 
ory, Equipment, and Methods. World Min., 
V. 32, No. 12, Nov. 1979, pp. 44-51. 

18. Leonard, J. W. , and D. R. Mitchell 
(ed.). Coal Preparation. AIME, 3d ed. , 
1968, 877 pp. 

19. Mcllvaine Co. (Northbrook, IL) . 
The Liquid Filtration Manual. 1980, pp. 
451-464. 

20. Moos, S. M. , and R. E, Dugger, 
Vacuum Filtration: Available Equipment 
and Recent Innovations. Min. Eng., v. 
31, No. 10, Oct, 1979, pp. 1473-1486. 

21. Parkson Corp. Fort Lauderdale, 
FL). Lamella Gravity Set tiers/ Thicken- 
ers. Bull. LT-103, 1979, 10 pp. 

22. Perrin, W. R. Co. , Ltd. (Houston, 
TX) . An Introduction to Filter Presses 
for Effluent and Sludge Dewatering. Un- 
dated, 16 pp. 



48 



23. Perry, R. H. , and C. H. Chilton 
(eds.). Chemical Engineers Handbook. 
McGraw-Hill, 5th ed. , 1973, 1958 pp. 

24. Pollution Equipment News. Select- 
ing Sludge Thickening and Dewatering 
Equipment. V. 13, No. 5, Oct. 1980, 
pp. 78-82. 

25. Schlitt, W. J., B. P. Ream, L. J. 
Haug, and W. D. Southard. Precipitating 
and Drying Cement Copper at Kennecott's 
Bingham Canyon Facility. Min. Eng. , 
V. 31, No. 6, June 1979, pp. 671-678. 

26. Schlitter, W. E., and W. Markl. 
Cross-Flow Lamella Thickeners. Min. 
Mag., V. 134, No. 4, Apr. 1976, pp. 261- 
297. 

27. Smelley, A. G. , and I. L. Feld. 
Flocculation Dewatering of Florida Phos- 
phatic Clay Wastes. BuMines RI 8349, 
1979, 26 pp. 

28. Soderberg, R. , and K. R. Dorman. 
Sludge Dewatering by Belt Press. Min. 
Cong. J., V. 65, No. 8, Aug. 1979, 
pp. 29-32. 

29. Sperry, D. R. and Co. (North 
Aurora, XL). Sperry Filter Presses. 
Catalog 12, undated, 29 pp. 

30. Sprute, R. H. , and D. J. Kelsh. 
Electrokinetic Densif ication of Slimes. 



Pres. at Northwest Mining Assoc. Ann. 
Conv. , Spokane, WA, Nov. 30-Dec. 1, 1978, 
25 pp.; available from the authors at the 
Bureau of Mines, Spokane Research Center, 
Spokane, WA. 

31. Star Systems, Filtration Division 
(Timmonsville, SC). Round and Square 
Filter Presses. 1979, 15 pp. 

32. Svarovsky, L. (ed.). Solid-Liquid 
Separation. Butterworths , 1979, 333 pp. 

33. Taggert, A. F. (ed.). Handbook of 
Mineral Dressing. 1947, 1915 pp. 

34. Thomas Publishing Co. 1981 Thomas 
Register of American Manufacturers and 
Thomas Register Catalogue File. 1981, 
8280 pp. 

35. Thrush, P. W. , and Staff, Bureau 
of Mines. A Dictionary of Mining, Miner- 
al, and Related Terms. BuMines, 1968, 
1269 pp. 

36. Wilson, E. B. , and F. G. Miller. 
Coal Dewatering — Some Technical and Ec- 
onomic Considerations. Min. Cong. J., 
V. 60, No. 9, Sept. 1974, pp. 116-121. 

37. Zatko, J. R. , B. J. Scheiner, 
and A. G. Smelley, Preliminary Studies 
on the Dewatering of Coal-Clay Waste 
Slurries Using a Flocculant. BuMines 
RI 8636, 1982, 15 pp. 



49 



APPENDIX A. —MATHEMATICAL TERMS 

Ac = The centrifugal acceleration of a particle in a centrifuge. 

Acs = The cross-sectional area of a gravitational settling basin. 

Aj = The drying area of a thermal dryer. 

A I = The inclination angle above horizontal of the plates in a multiple-plate 
thickener. 

Aq = The area of the orifices in the constriction plate of a fluidized-bed dryer. 

Ap = The top surface area of each plate in a multiple-plate thickener. 

At I = The area of a vacuum filter test leaf. 

Cave - The average solids concentration in the compression zone of a gravitational 
thickener. 

C(j = The discharge coefficient that characterizes the airflow through an orifice or 
constriction plate of a fluidized-bed dryer. 

Cs2 = The secondary solids concentration produced when a hypothetical gravitational 
thickener processes a slurry having the maximum solids concentration permis- 
sible for that particular thickener, 

Cse - The solids concentration in a settling test after an arbitrary time, t^. 

Css - The solids concentration of the slurry prior to dewatering or thickening. 

Cssi = The solids concentration of a hypothetical slurry as shown on a batch flux 
versus solids concentration graph. 

Css2 ~ The solids concentration of a hypothetical slurry that is greater than Cggi. 

Csu = The solids concentration of the underflow or thickened solids. 

Csui = The solids concentration of the underflow produced by a hypothetical gravita- 
tional thickener processing an arbitrary slurry having a solids concentration 
CssI • 

Csu2 ~ The solids concentration of a hypothetical underflow that is greater than 
Csu I • 

D| = The density of the slurry liquid at a specified temperature. 

Dm = The diffusivity value of moisture through a solid during thermal drying. 

Ds = The density of the solid particles in a slurry, 

Ec = The rate of evaporation during the constant-rate drying period during thermal 
drying. 

Ef = The evaporation rate during the falling rate period during thermal drying. 



50 

Ef = The total efficiency for a particular piece of solid-liquid separation 
equipment. 

e = Euler's number, which has a value of 2.71828... 

Fbt = The bulk transport flux component based on the underflow from a particular 
gravitational thickener. 

Fc = Centrifugal force expressed as multiples of gravitational force. 

Fg = The settling flux component for a particular gravitational thickener. 

Fg I = An intercept point on the batch flux axis of a hypothetical batch flux versus 
solids concentration graph for a gravitational thickener. 

F-i-s = The total solids flux for a particular gravitational thickener. 

Fy = The filter cake correlating factor for vacuum filters. 

fcv = The filter cake formation rate for a vacuum filter. 

fcvn - The iiew cake formation rate for a vacuum filter. 

fcvo ~ The old cake formation rate for a vacuum filter. 

G = The gravitational acceleration at the earth's surface. 

G(x) = A gravimetric separation function that describes the separation efficiency of 
a piece of solid-liquid separation equipment. 

Hg = The humidity of the ambient air around a thermal dryer. 

Ha be - The height of the interface between zones A and B in a slurry settling test 
after an arbitrary time, tg. 

Hgbo - The original height of the interface between zones A and B in a slurry set- 
tling test. 

Hg I = The humidity of the air at the solid-air interface in a thermal dryer. 

hy = A heat transfer coefficient for thermal drying. 

J = A correction factor for particle shape. 

Ky = A mass transfer coefficient for thermal drying. 

Lp = The length of each plate in a multiple-plate thickener. 

Ls2 = The length of the settling zone in the direction of slurry transport in a 
centrifuge. 

Mg = The molecular weight of air. 

Mg = The mass of all solids that have been separated from the liquid of a slurry. 



51 

Mf = The total mass of all solids in a slurry before solid-liquid separation. 

N = The number of revolutions per second for a centrifuge bowl. 

P(jc = The pressure differential across the cake of a vacuum filter. 

Pdo = The air pressure differential across the orifice plate of fluidized-bed 
dryers. 

Pop = The mean absolute air pressure in the orifice of a constriction plate in a 
fluidized-bed dryer. 

Qc = The volumetric flow of slurry through a centrifuge. 

Qs = The volumetric flow of a slurry through a settling zone. 

qa = The airflow through the orifice plate under controlled air pressure and tem- 
perature conditions in a fluidized-bed dryer. 

Re = The radius of curvature of a particle in a centrifuge. 

Rcb = The radial distance from the rotational axis to the inside surface of the bowl 
in a centrifuge. 

Rsc = The percent recovery of solids in the cake produced by a centrifuge. 

Rsz = The radial distance from the rotational axis to the surface of the settling 
zone in a centrifuge. 

S = Equal to one-half the thickness of the layer of solids in a thermal dryer. 

Sf = The subsidence rate for solids in the bottom of a settling tank. 

Sgs = The average specific gravity of the solids in a slurry. 

Sgsc ~ The average specific gravity of the solids in compression. 

Sg,y = The specific gravity of water, which is 1.0. 

Sp = The spacing between plates of a multiple-plate thickener. 

Tg = The ambient air temperature around a thermal dryer. 

Ti = The air temperature at the solid-air interface in a thermal dryer. 

Tk = The temperature in Kelvins. 

tcv = The cycle time for vacuum filters. 

tcvn - The new cycle time for a vacuum filter. 

tcvo = The old cycle time for vacuum filter, 

t(jv = The drying time for a vacuum filter cake. 



52 

te = An arbitrary amount of time after the beginning of a settling test for gravi- 
tational thickeners. 

ths = The holding time necessary for the solids to settle from the entering feed 
solids concentration to the underflow concentration, 

tn = An arbitrary time after the falling-rate drying has begun in a thermal dryer, 

Vam = The air mass velocity through a thermal dryer, 

Vav = The volume of air passing through a vacuum filter cake per unit of filter sur- 
face area, 

Vsc - The volume of solids in compression in a gravitational thickener, 

Vjjt = The bulk transport velocity of the solids through a gravitational thickener. 

V| = The linear velocity of a slurry particle moving through a settling zone. 

Vp = The linear peripheral velocity of a particle in a centrifuge, 

Vg = The terminal settling velocity of solids through a liquid, 

Vy = The resultant velocity vector of a settling slurry particle as a consequence 
of its forward motion and the downward pull of gravity. 

The dry weight of the solids cake obtained during a vacuum filter leaf test. 

The dry weight of solids for a given cake thickness on a vacuum filter. 

The weight ratio of water to solids in the discharge of a gravitational 
thickener. 

The weight ratio of water to solids in the slurry. 

The weight percent of solids in the cake. 

The weight percent of solids in the liquid after liquid-solid separation. 

The weight percent of solids in a slurry prior to liquid-solid separation. 

The equilibrium moisture content of the solids during the falling-rate period 
of thermal drying, 

Xfav = The average free moisture content during the falling-rate period of thermal 
drying. 

Xf i = The initial free moisture content of the solids at the beginning of the 
falling-rate period during thermal drying. 

Xfav = The average total solids moisture content during the falling-rate period of 
thermal drying, 

Xt i = The initial total moisture content of the solids at the beginning of the 
falling-rate period of thermal drying. 



Wds 


Wdst 


Wrd 


Wrs 


Wsc 


Wsl 


Wss 


Xen 



53 



X = The diameter of solid particles in a slurry, 

Xmax = The largest particle size on a grade efficiency curve that is capable of being 
separated from the liquid of a slurry by a particular piece of dewatering 
equipment. 

X50 = The particle size on a grade efficiency curve representing a 50-pct probabil- 
ity of being separated from the liquid of a slurry. 

X98 = The particle size on a grade efficiency curve that represents 98-pct separa- 
tion efficiency. 

a = A factor used in the liquid diffusion equation for the falling-rate drying 
period of thermal drying. 

g = A factor used in the liquid diffusion equation for the falling-rate drying 
period of thermal drying. 

Xj = The latent heat of water at the temperature of the solid-air interface in a 
thermal dryer, 

y = The viscosity of the liquid in a slurry at a specified temperature, 

TT = 3.1416. 

Z = A characteristic value that describes the machine variables for a particular 
centrifuge. 

0) = The angular velocity of a particle undergoing centrifugal acceleration. 



54 



APPENDIX B. —MANUFACTURERS OF DEWATERING EQUIPMENT 
AS OF OCTOBER 1982 (24. 34)1.2 



AFL Industries, Inc. 

3661 West Blue Heron Blvd. 

Riviera Beach, FL 33404 

Al f a-Laval , Inc . 
2115 Linwood Ave, 
Fort Lee, NJ 07024 

Baker-Perkins, Inc. 
1000 Hess St. 
Saginaw, MI 48601 

The Leon J. Barrett Co. 

Box 551 

Worcester, MA 01613 

Bird Machine Co. , Inc. 
100-T Neponset St. 
South Walpole, MA 02071 

Calgon Corp. 
Box 1346-C 
Pittsburgh, PA 15230 

Carus Chemical Co. , Inc. 
1500 Eighth St. 
LaSalle, IL 61301 

C-E Bauer 

Box 968 

Springfield, OH 45501 

C-E Raymond 

200 West Monroe St. 

Chicago, IL 60606 



Dedert Corp. 

20000-T Governor's Dr. 

Olympia Field, IL 60461 

Denver Equipment Div. 
Joy Manufacturing Co. 
Box 340 
Colorado Springs, CO 80901 

Donaldson Co. 
Liquid Systems Div. 
1400 West 94th St. 
Minneapolis, MN 55431 

Dorr'-Oliver, Inc. 
79 Havemeyer Lane 
Stamford, CT 06904 

Duriron Co. , Inc. 
Filtration Systems Div. 
9542 Hardpan Rd. 
Angola, NY 14006 

Environmental Elements Corp. 
Box 1318 
Baltimore, MD 21203 

Envirotech Corp. 
3000 Sand Hill Rd. 
Menlo Park, CA 94025 

ERC/Lancy Div. , Dart 

& Kraft Co. 
525 West New Castle St. 
Zelienople, PA 16063 



Centrifugal and Mechanical 

Industries, Inc. 
146 President St. 
St. Louis, MO 63118 

Clow Corp. 
1211 W. 22d St. 
Oak Brook, IL 60521 



Filpaco Industries, Inc. 
3837 West Lake St. 
Chicago, IL 60624 

FMC Corp., Materials 

Handling Systems Div. 
3400 Walnut 
Colmar, PA 18915 



1 Reference to specific equipment suppliers does not imply endorsement by the Bureau 
of Mines. 

^This list is as complete as possible based on the information available at the 
time this paper was written. No responsibility can be taken for omissions or changes 
in listings. 



55 



Heyl and Patterson 

Dept. 10 

Box 36 

Pittsburgh, PA 15230 

Industrial Filter & Pump 
Manufacturing 
5900 West Ogden Ave. 
Cicero, XL 60650 

Infilco Degremont, Inc. 
Dept. T-R 
Box 29599 
Richmond, VA 23288 

lU Conversion Systems , Inc. 
Dept. T-R 
115 Gibraltar Rd. 
Horsham, PA 19044 

JWI, Inc. 

Box 9A 

Holland, MI 49423 

Keene Corp. Filtration Div. 
1571 Forrest Ave. 
LaGrange, GA 37743 

Komline-Sanderson Engineering 
Corp. 

100 Holland Ave. 
Peapack, NJ 07977 

Krebs Engineers 
1205 Chrysler Dr. 
Menlo Park, CA 94025 

Krofta Engineering Corp. 
101-T Yokun Ave. 
Lenox, MA 01240 

Lakos Separators 
1911 North Helm 
Box 6119 
Fresno, CA 93703 

Larox OY 

Box 29 

SF-53101 Lappeenranta 10 

Finland 



Lavin/Guinard International, 

Inc. 
500 Davisville Rd. 
Hatboro, PA 19040 

McNally-Pittsburg 

Manufacturing Corp. 
Third at Walnut St. 
Pittsburg, KS 66762 

Nalco Chemical Co. 
2901 Butterfield Rd. 
Oak Brook, IL 60521 

National-Standard Co. 
Perforated Metals Div. 
Drawer 507 
Carbondale, PA 18407 

Netzsch, Inc. 
119 Pickering Way 
Pickering Creek Industrial 

Park 
Ext on, PA 19341 

Parkson Corp. 

2727-T NW 62d St. 

Fort Lauderdale, FL 33309 

Passavant Corp, 
Carson Rd. 
Box 2503 
Birmingham, AL 35201 

William R. Perrin Co., Ltd. 
432 Monarch Ave. 
Aj ax , Ont . 
Canada LIS 2G7 

Serfilco Div. 

Service Filtration Corporation 

1234 Depot St. 

Glenview, IL 60025 

SFS Div. , BINAB USA, Inc. 
15271 NW 60th Ave. 
Miami Lakes, FL 33014 

D. R. Sperry and Co. 
112-T North Grant St. 
North Aurora, IL 60542 



56 



Star Systems Filtration Div. 

101 Kershaw St. 

Box 815 

Timmonsville, SC 29161 

Transamerica Delaval, Inc. 
Condenser and Filter Div. 
Front St. 
Florence, NJ 08518 

Tretolite Div. , Tretolite Corp. 
369 Marshall Ave. 
St. Louis, MO 63119 

Vara International, Inc. 
1201-T 19th PI. 
Vara International Plaza 
Vero Beach, FL 32960 



WEMCO Div., Envirotech Corp, 
1796 Tribute Rd. 
Box 15619 
Sacramento, CA 95813 

Western States Machine Co. 
1716 Fair grove Ave. 
Hamilton, OH 45012 

Zimpro, Inc. 
Dept. MZ 
Military Road 
Rothschild, WI 54474 



57 



APPENDIX C. —AVAILABLE DEWATERING EQUIPMENT LISTED BY MANUFACTURER 

AS OF OCTOBER 1982 (24) 



Equipment and manufacturer^ »^ 
Gravitational thickeners: 

AFL 

Denver 

Dorr-Oliver 

Envlrotech 

Industrial Filter & Pump 

Larox 

Parkson 

SFS 

Centrifuges: 

Alf a-Laval. 

Baker-Perkins 

Barrett 

Bird Machine 

Centrifugal and Mechanical... 

Dedert 

Donaldson 

Dorr-Oliver 

Envlrotech 

lU Conversion 

Lavln 

MET Pro 

National Standard 

WEMCO 

Western States 

Hydrocyclones : 

C-E Bauer 

Dorr-Oliver 

Krebs Engineers 

Lakos 

Larox 

WEMCO 

Filters: 

AFL 

Bird Machine 

C-E Bauer 

Clow 

Denver 

Dorr-Oliver 

Durlron 

Environmental Elements 

Envlrotech 

ERC/Lancy 

Fllpaco 

Industrial Filter & Pump 

Inf llco 

lU Conversion 

JWI 

Komllne-Sanderson 

Krofta 

See footnotes at end of table. 



Type 



Multiple plate. 

Conventional. 

Conventional and high rate. 

Do. 
Multiple plate. 
Conventional. 
Multiple plate. 

Do. 

Basket, disk, and solid bowl. 

Pusher. 

Basket and solid bowl. 

Do. 
Basket. 

Do. 
Solid bowl. 

Basket, disk, and solid bowl. 
Horizontal. 
Solid bowl. 

Basket and solid bowl. 
Solid bowl. 
Basket. 

Do. 

Do. 

Undifferentiated. 
Do. 
Do. 
Do. 
Do. 
Do. 

Belt press and gravity. 

Horizontal belt, belt, and drum. 

Gravity and rotary. 

Filter press and horizontal belt. 

Horizontal belt. 

Belt, disk, and drum. 

Filter press. 

Do. 
Drum, filter press, gravity, and horizontal belt, 
Filter press. 

Do. 

Do. 
Horizontal belt. 
Drum. 

Filter press. 

Vacuum belt, belt press, drum, and gravity. 
Belt and gravity. 



58 



APPENDIX C. —AVAILABLE DEWATERING EQUIPMENT LISTED BY MANUFACTURER 
AS OF OCTOBER 1982 (24)— Continued 



Equipment and manufacturer^ '^ 



Type 



Filters — Continued: 

Larox 

Netzsch Inc 

Parkson 

Passavant 

Perrin. 

Serfilco 

Sperry 

Star Systems 

Transamerica-Delaval 

Vara 

Zimpro 

Thermal dryer or incinerator; 

C-E Raymond 

Dedert. 

Dorr-Oliver 

FMC , 

Heyl and Patterson , 

Komline- Sanderson , 

McNally-Pittsburg , 

Chemical treatment aids: 

Calgon 

Carus ( 

ERC/Lancy , 

Industrial Filter & Pump.., 

lU Conversion , 

Keene , 

Nalco , 

Passavant , 

Tretolite , 



' See appendix B for comple 
^This list is as complete 

time this paper was written, 

in listings. 



Disk, filter press, and drum. 

Filter press. 

Belt press. 

Belt press and filter press. 

Horizontal belt and filter press. 

Gravity and vacuum belt. 

Filter press. 

Do. 
Gravity and belt press. 
Gravity. 
Filter press. 

Rotary, fluidized bed, and suspension. 

Drum and suspension. 

Fluidized bed. 

Drum, fluidized bed, and louvre. 

Fluidized bed. 

Suspension. 

Fluidized bed and vertical tray. 

Polyelectroly tes . 

Potassium permanganate. 

Lime and polyelectrolytes. 

Lime. 

Lime and polyelectrolytes. 

Do. 
Polyelectrolytes . 
Lime and polyelectrolytes. 
Polyelectrolytes . 



te name and address, 
as possible based on the information available at the 
No responsibility can be taken for omissions or changes 



59 



APPENDIX D. —EQUIPMENT EFFICIENCY 



Less than perfect performance of sep- 
aration equipment can be characterized by 
the separation efficiency. The grade ef- 
ficiency concept can be applied to solid- 
liquid separation equipment whose per- 
formance does not change with time if all 
operational variables are kept constant. 
Hydrocyclones , centrifuges, and gravita- 
tional thickeners are examples of such 
equipment. This concept is not widely 
used in filtration because the efficiency 
changes with the amount of solids col- 
lected on the face of the filter medium. 
For filtration, though, it is helpful to 
determine the grade efficiency of the 
clean medixom, which influences the ini- 
tial retention characteristics of the 
filter and can be used for filter rating 
(32, p. 31).'' 

TOTAL EFFICIENCY 

The total efficiency for dewater- 
ing equipment can be determined by the 
equation 



M< 
Mt' 



(D-1) 



Additional information must be known 
about the particle-size distribution of 
the feed solids , the density of the sol- 
ids , and such operational data as flow 
rate, temperature, type of fluid, and 
solids input concentration. A single 
value for the total efficiency cannot be 
used to represent the separation capabil- 
ity of the equipment for any materials 
other than those actually tested ( 32 , 
pp. 34-35). 

GRADE EFFICIENCY 

The efficiency of separation equipment, 
however, can be characterized by a gravi- 
metric grade efficiency function, G(x) . 
This is a probablistic mathematical ex- 
pression, based on mass efficiency, which 
describes the particle trajectories dur- 
ing the separation process. A grade ef- 
ficiency function can be developed for 
each type of separation equipment that 
describes the efficiency of separation 
for a range of particle sizes. This 
information can be graphed as an S-shaped 
curve, such as the one shown in fig- 
ure D-1. This graph is often referred to 



where Ef = the total equipment effi- 
ciency, 

Mg = the mass of all solids sepa- 
rated from a slurry liquid, 

and Mf = the mass of all slurry sol- 
ids prior to solid-liquid 
separation (32, p. 33). 

The performance of most available sepa- 
rational equipment is predominantly size 
dependent, so the total efficiency de- 
pends on the size distribution of the 
feed solids and is not suitable for the 
general criterion of efficiency. Conse- 
quently values of total efficiency stated 
by equipment manufacturers may not be en- 
tirely accurate concerning the separa- 
tional capability of their equipment. 

'Underlined numbers in parentheses re- 
fer to items in the list of references 
preceding the appendixes. 



u 
a 

>■" 

o 

z 

UJ 

o 

u. 
u. 

Ui 

z 

o 

H 
< 

< 
Q. 
UJ 
(0 



QC 

< 
Q. 



too 

98 



80 

60 
50 
40 



KEY 

Particle size having 
5 0-p ct separation 
efficiency 
— X Particle size having 

9^ 90T3ct 

separation efficiency/ 
Particle size having 
^ 10 0-pct 

separation 
efficiency 



_i 20 




50 



PARTICLE SIZE,(dimensionless) 

FIGURE D-1. - Example of a grade efficiency 
curve showing the relationship of Xjq, X93, and 
X.., (32). 



60 



as the partition probability curve be- 
cause it shows the probability for each 
particle size of either being separated 
or remaining with the fluid ( 32 , p. 35). 

PARTICLE SIZE PARAMETERS 

The grade efficiency curve can be used 
for determining several operational pa- 
rameters for a particular piece of equip- 
ment. The particle corresponding to the 
50-pct probability is called the equi- 
probable size, X30, and is used as the 
minimum cutoff size or cut size of the 
particular type of equipment. This cut 
size is independent of the feed material, 
and its determination requires a knowl- 
edge of the entire grade efficiency curve 
(32, p. 35). 

In any separation operation, there will 
be a particle size larger than the grade 
efficiency, which is 100 pet. This is 
the largest particle remaining in the 
overflow after separation of the maximum 



particle size that would have a chance to 
escape and is called x^ax ( 32 , p. 38). 

If the particle trajectories in the 
separator can be approximated, the most 
unfavorable conditions of separation are 
taken for detenaining this limit of sepa- 
ration. It is difficult to determine the 
limit of separation accurately, so the 
size corresponding to 98-pct efficiency, 
X98, is used, which gives an easily de- 
fined point. This size is called the 
approximate limit of separation and is 
widely used in filter rating ( 32 , p. 38). 
The relationship of X50, xgs, and x^ax is 
shown in figure D-1. 

The concept of grade efficiency is 
helpful in determining the application of 
a particular piece of equipment for a 
particular dewatering or desliming opera- 
tion. For a more detailed discussion of 
grade efficiency, consult Svarovsky ( 32 , 
pp. 31-57). 



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